Chlorine
1
Chlorine
Peter Schmittinger, CREANOVA GmbH, Werk Lülsdorf, Niederkassel, Federal Republic of Germany
(Chap. 1, 2, 3, 4, 5, 9, 10, 11, 12, 13, 14 and 15)
Thomas Florkiewicz, OxyTech Systems, Chardon, Ohio, United States (Chap. 6 and 9)
L. Calvert Curlin, OxyTech Systems, Chardon, Ohio, United States (Chap. 6 and 9)
Benno Lüke, Uhde, Dortmund, Federal Republic of Germany (Chap. 7 and 9)
Robert Scannell, DeNora Deutschland GmbH, Rodenbach, Federal Republic of Germany (Chap. 8.1)
Thomas Navin, OxyTech Systems Inc., Chardon, Ohio, United States (Chap. 8.2)
Erich Zelfel, Infraserv, Knapsack, Federal Republic of Germany (Chap. 12)
Rüdiger Bartsch, Technische Universität München, München, Federal Republic of Germany (Chap. 16)
1.
2.
3.
4.
4.1.
4.2.
5.
5.1.
5.2.
5.2.1.
5.2.2.
5.2.3.
5.2.4.
5.3.
5.3.1.
5.3.2.
5.3.3.
5.3.4.
5.3.5.
6.
6.1.
6.2.
6.2.1.
6.2.2.
6.2.3.
6.2.4.
6.2.5.
6.3.
6.3.1.
6.3.2.
6.3.3.
6.3.4.
6.3.5.
7.
7.1.
7.2.
Introduction . . . . . . . .
Physical Properties . . . .
Chemical Properties . . .
Chlor-Alkali Process . .
Brine Supply . . . . . . . .
Electricity Supply . . . .
Mercury Cell Process . .
Principles . . . . . . . . . .
Mercury Cells . . . . . . .
Uhde Cell . . . . . . . . . .
De Nora Cell . . . . . . . .
Olin – Mathieson Cell . . .
Solvay Cell . . . . . . . . .
Operation . . . . . . . . . .
Brine System . . . . . . . .
Cell Room . . . . . . . . . .
Treatment of the Products
Measurement . . . . . . . .
Mercury Emissions . . . .
Diaphragm Process . . .
Principles . . . . . . . . . .
Diaphragm Cells . . . . .
Dow Cell . . . . . . . . . .
Glanor Electrolyzer . . .
OxyTech “Hooker” Cells
HU Monopolar Cells . .
OxyTech MDC Cells . .
Operation . . . . . . . . . .
Brine System . . . . . . . .
Cell Room . . . . . . . . . .
Diaphragm Aging . . . . .
Treatment of the Products
Measurement . . . . . . . .
Membrane Process . . . .
Principles . . . . . . . . . .
Process Specific Aspects
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2
3
6
8
11
13
13
14
19
19
20
21
21
21
21
23
24
24
25
27
28
31
33
34
35
35
37
37
37
39
41
41
43
43
44
46
c 2006 Wiley-VCH Verlag GmbH & Co. KGaA, Weinheim
10.1002/14356007.a06 399.pub2
7.2.1.
7.2.2.
7.2.3.
7.2.4.
7.3.
7.3.1.
7.3.2.
7.3.3.
7.3.4.
8.
8.1.
8.1.1.
8.1.2.
8.1.3.
8.1.4.
8.2.
9.
9.1.
9.2.
9.2.1.
9.2.2.
9.2.3.
10.
10.1.
10.2.
10.2.1.
10.2.2.
10.2.3.
11.
11.1.
11.2.
11.3.
Brine Purification . . . . . . . . . . .
Commercial Membranes . . . . . . .
Power Consumption . . . . . . . . . .
Product Quality . . . . . . . . . . . .
Membrane Cells . . . . . . . . . . .
Monopolar and Bipolar Designs . .
Commercial Electrolyzers . . . . . .
Comparison of Electrolyzers . . . .
Cell Room . . . . . . . . . . . . . . . .
Electrodes . . . . . . . . . . . . . . .
Anodes . . . . . . . . . . . . . . . . . .
General Properties of the Anodes .
Anodes for Mercury Cells . . . . . .
Anodes for Diaphragm Cells . . . .
Anodes for Membrane Cells . . . . .
Activated Cathode Coatings . . . .
Comparison of the Processes . . .
Product Quality . . . . . . . . . . . .
Economics . . . . . . . . . . . . . . .
Equipment . . . . . . . . . . . . . . . .
Operating Costs . . . . . . . . . . . .
Summary . . . . . . . . . . . . . . . .
Other Production Processes . . . .
Electrolysis of Hydrochloric Acid
Chemical Processes . . . . . . . . .
Catalytic Oxidation of Hydrogen
Chloride by Oxygen . . . . . . . . . .
Oxidation of Hydrogen Chloride by
Nitric Acid . . . . . . . . . . . . . . .
Production of Chlorine from Chlorides . . . . . . . . . . . . . . . . . . .
Chlorine Purification and Liquefaction . . . . . . . . . . . . . . . . . .
Cooling . . . . . . . . . . . . . . . . .
Chlorine Purification . . . . . . . .
Drying . . . . . . . . . . . . . . . . . .
47
50
51
51
51
51
53
57
57
57
57
59
60
61
62
62
63
64
64
64
66
66
67
67
69
69
71
71
71
71
72
72
2
11.4.
11.5.
11.6.
12.
12.1.
12.2.
12.3.
12.4.
12.5.
12.6.
Chlorine
Transfer and Compression . . .
Liquefaction . . . . . . . . . . . . .
Chlorine Recovery . . . . . . . . .
Chlorine Handling . . . . . . . . .
Storage Systems . . . . . . . . . .
Transport . . . . . . . . . . . . . . .
Chlorine Discharge Systems . . .
Chlorine Vaporization . . . . . .
Treatment of Gaseous Effluents
Materials . . . . . . . . . . . . . . .
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73
74
75
76
76
77
78
79
80
80
1. Introduction
Although C. W. Scheele reported the formation
of chlorine gas from the reaction of manganese
dioxide with hydrochloric acid in 1774, he did
not recognize the gas as an element [37]. H. Davy
is usually accepted as the discoverer (1808),
and he named the gas chlorine from the Greek
κλω̃oσ (chloros), meaning greenish yellow.
Chlorine for bleaching textiles was first produced from manganese dioxide and hydrochloric acid by a process developed by Weldon,
the yield of chlorine being 35 % of the theoretical value. In 1866, Deacon developed a process based on the oxidation of hydrogen chloride
gas by atmospheric oxygen in the presence of a
copper salt, CuCl2 , as the catalyst and obtained
yields up to 65 % of the theoretical value.
In 1800, Cruickshank was the first to prepare chlorine electrochemically [38]; however,
the process was of little significance until the
development of a suitable generator by Siemens
and of synthetic graphite for anodes by Acheson and Castner in 1892. These two developments made possible the electrolytic production
of chlorine, the chlor-alkali process, on an industrial scale. About the same time, both the diaphragm cell process (1885) and the mercury
cell process (1892) were introduced. The membrane cell process was developed much more
recently (1970). Currently, more than 95 % of
world chlorine production is obtained by the
chlor-alkali process. Since 1970 graphite anodes
have been superseded by activated titanium anodes in the diaphragm and mercury cell processes. The newer membrane cell process uses
only activated titanium anodes.
Other electrochemical processes in which
chlorine is produced include the electrolysis of
hydrochloric acid and the electrolysis of molten
12.7.
13.
13.1.
13.2.
14.
15.
16.
17.
18.
Safety . . . . . . . . . . . . . . . . . .
Quality Specifications and Analytical Methods . . . . . . . . . . . . . .
Quality Specifications . . . . . . . .
Analytical Methods . . . . . . . . .
Uses . . . . . . . . . . . . . . . . . . .
Economic Aspects . . . . . . . . . .
Toxicology . . . . . . . . . . . . . . .
Acknowledgement . . . . . . . . . .
References . . . . . . . . . . . . . . .
81
82
82
82
83
86
87
89
89
alkali metal and alkaline earth metal chlorides, in
which the chlorine is a byproduct. Purely chemical methods of chlorine production are currently
insignificant.
Since 1975, the membrane cell process has
been developed to a high degree of sophistication. It has ecological advantages over the
mercury processes and has become the most
economically advantageous process. The membrane cell process has become widely accepted,
and all new plants are using this technology. By
2000 more than 30 % of the chlorine will be produced in membrane cells.
World capacity for chlorine exceeds
45 × 106 t/a. With an annual energy consumption of about 1.5 × 1011 kWh, the chlor-alkali
process is one of the largest industrial consumers
of electrical energy. The chlorine production of
a country is an indicator of the state of development of its chemical industry.
Occurrence and Formation. Chlorine is
the 11th most abundant element in the lithosphere. Because it is highly reactive, it is rarely
found in the free state and then mainly in volcanic gases. It exists mainly in the form of
chlorides, as in sea water, which contains an average of 2.9 wt % sodium chloride and 0.3 wt %
magnesium chloride. In salt deposits formed by
evaporation of seas, there are large quantities of
rock salt (NaCl) and sylvite (KCl), together with
bischofite (MgCl2 · 6 H2 O), carnallite (KCl ·
MgCl2 · 6 H2 O), tachhydrite (CaCl2 · 2 MgCl2
· 12 H2 O), kainite (KCl · MgSO4 · 3 H2 O), and
others. Occasionally there are also heavy metal
chlorides, usually in the form of double salts,
such as atacamite (CuCl2 · 3 Cu(OH)2 ), and
compounds of lead, iron, manganese, mercury,
or silver. Chlorates and perchlorates occur to a
small extent in Chile saltpeter. Free hydrochloric
Chlorine
acid is occasionally found in gases and springs
of volcanic origin. Plants and animals always
contain chlorine in the form of chlorides or free
hydrochloric acid.
Chlorine is formed by oxidation of hydrochloric acid or chlorides by such compounds
as manganese dioxide, permanganates, dichromates, chlorates, bleaching powder, nitric acid,
or nitrogen oxides. Oxygen, including atmospheric oxygen, acts as an oxidizing agent in
the presence of catalysts. Some metal chlorides produce chlorine when heated, for example, gold(III) chloride or platinum chloride.
3
The density of chlorine gas at 101.3 kPa is a
function of temperature:
t, ◦ C
, kg/m3
0
3.213
50
2.700
100
2.330
150
2.051
The density up to 300 ◦ C is higher than that
of an ideal gas because of the existence of more
complex molecules, for example, Cl4 . In the
range 400 – 1450 ◦ C, the density approximates
that of an ideal gas, and above 1450 ◦ C thermal dissociation takes place, reaching 50 % at
2250 ◦ C. The density of chlorine gas as a function of temperature and pressure is shown in Figure 1. The gas state can be described by the van
der Waals equation with
2. Physical Properties
Chlorine [7782-50-5], EINECS no. 231–959–5,
exists in all three physical states. At STP it is a
greenish-yellow pungent, poisonous gas, which
liquefies to a mobile yellow liquid. Solid chlorine forms pale yellow rhombic crystals. The
principal properties are given below; more details, including thermodynamic values are given
in [40] and in “New Property Tables of Chlorine
in SI Units” [41].
Atomic number Z
Relative atomic mass Ar
Stable isotopes (abundance)
Electronic configuration in
the ground state
Term symbol in the ground
state
Melting point mp
Boiling point bp
Critical density crit
Critical temperature T crit (t crit )
Critical pressure pcrit
Density of gas (0 ◦ C, 101.3 kPa)
Density relative to air d
Enthalpy of fusion ∆H f
Enthalpy of vaporization ∆H v
Standard electrode potential E ◦
Enthalpy of dissociation ∆H diss
Electron affinity A
Enthalpy of hydration ∆H hyd
of Cl−
Ionization energies ∆E i
EC No.
a = 6.580 L2 bar mol−2 , b = 0.05622 L/mol
17
35.453
35 (75.53 %)
37 (24.47 %)
[Ne] 3s2 3p5
2
P3/2
172.17 K (− 100.98 ◦ C)
239.10 K (− 34.05 ◦ C)
565.00 kg/m3
417.15 K (144.0 ◦ C)
7.71083 MPa
3.213 kg/m3
2.48
90.33 kJ/kg
287.1 kJ/kg
1.359 V
239.44 kJ/mol
(2.481 eV)
364.25 kJ/mol (3.77 eV)
405.7 kJ/mol
13.01, 23.80, 39.9, 53.3,
67.8, 96.6, 114.2 eV
017–001–00–7
Figure 1. Density of chlorine gas as a function of
temperature and pressure
The density of liquid chlorine is given in
Figure 2. The compressibility of liquid chlorine is the greatest of all the elements. The
volume coefficient per MPa at 20 ◦ C over the
range 0 – 10 MPa is 0.012 %. The coefficient
increases rapidly with temperature: 0.023 % at
35 ◦ C, 0.037 % at 64 ◦ C, and 0.064 % at 91 ◦ C.
One liter of liquid chlorine at 0 ◦ C produces
456.8 dm3 of chlorine gas at STP; 1 kg of liquid produces 311 dm3 of gas.
4
Chlorine
Thermodynamic information is given in Table 1, from which the data required for working with gaseous and liquid chlorine can be obtained [42]. The Joule – Thomson coefficient is
0.0308 K/kPa at STP.
At STP the specific heats of chlorine are
cp = 0.481 kJ kg−1 K−1
cv = 0.357 kJ kg−1 K−1
κ = cp /cv = 1.347
The molar heat capacity at constant volume
cv increases with temperature [43]:
t, ◦ C
cv , J/mol
0
24.9
100
26.4
200
28.1
500
28.9
1000
29.7
Figure 2. Density of liquid chlorine
The vapor-pressure curve for chlorine is
shown in Figure 3.
The heat capacity of liquid chlorine decreases
over the temperature range − 90 ◦ C to 0 ◦ C:
t, ◦ C
− 90
−1
−1
c, J kg
K
c, J mol−1 K−1
− 70
− 50
− 30
0
0.9454 0.9404 0.9341 0.9270 0.9169
67.03 66.70 66.23 65.73 65.02
The thermal conductivities of chlorine gas
and liquid are almost linear functions of temperature from − 50 ◦ C to 150 ◦ C:
t, ◦ C
− 50
λg ,
6.08
W m−1 K−1 ×102
0.17
λl ,
W m−1 K−1
− 25
0
25
50
75
7.06
7.95
8.82
9.75
10.63 11.50
0.16
0.15
0.135 0.12
0.11
100
0.09
The viscosities of chlorine gas and liquid are
shown in Figure 4 over the same temperature
range. The surface tension at the liquid – gas interface falls rapidly with temperature:
Figure 3. Vapor pressure of liquid chlorine
t, ◦ C
− 50
− 25
0
25
50
σ, mJ/m2
29.4
25.2
20.9
16.9
13.4
The vapor pressure can be calculated over the
temperature range 172 – 417 K from the Martin – Shin – Kapoor equation [41]:
ln P = A+
B
E (F −T ) ln (F −T )
+ClnT +DT +
T
FT
A = 62.402508
B = − 4343.5240
C = − 7.8661534
D = 1.0666308×10−2
E = 95.248723
F = 424.90
Figure 4. Viscosities of chlorine gas and liquid
Chlorine
5
Table 1. Properties of liquid and gaseous chlorine [41]. Lower values are quoted in more recent literature [38, 39], especially in the region of
the critical points.
Temperature t,
◦
C
−70
−60
−50
−40
−30
−20
−10
0
10
20
30
40
50
60
70
80
90
100
110
120
130
140
144
Pressure, Specific volumes, dm3 /kg
p,
Specific enthalpies, kJ/kg*
Specific entropies, kJ kg−1 K−1
bar
liquid
vapor
liquid
vaporization
vapor
liquid
vaporization
vapor
0.1513
0.2768
0.4762
0.7772
1.212
1.816
2.628
3.689
5.043
6.731
8.800
11.30
14.27
17.76
21.84
26.55
31.95
38.14
45.18
53.18
62.24
72.50
77.01
0.6042
0.6135
0.6233
0.6336
0.6445
0.6560
0.6682
0.6812
0.6951
0.7100
0.7261
0.7435
0.7627
0.7837
0.8073
0.8339
0.8646
0,9010
0.9456
1.0039
1.0890
1.2624
1.7631
1563
894.4
541.8
344.9
229.0
157.7
112.1
81.89
61.26
46.77
36.35
28.66
22.88
18.44
14.97
12.20
9.944
8.082
6.508
5.169
4.001
2.842
1.763
351.11
360.69
370.15
379.70
389.37
399.21
408.88
418.68 **
428.43
438.19
447.90
457.66
467.45
477.50
487.76
498.56
510.25
523.35
537.88
554.62
575.10
603.74
642.30
306.89
301.58
296.29
290.73
284.95
278.84
272.73
266.28
259.67
252.80
245.72
238.31
230.53
222.07
212.90
202.60
190.79
176.85
160.14
139.59
113.30
71.18
0
658.00
662.27
666.41
670.43
674.33
678.05
681.61
684.96
688.10
690.99
693.63
695.97
697.98
699.57
700.66
701.16
701.04
700.20
698.02
694.21
688.39
674.91
642.30
3.9021
3.9481
3.9917
4.0336
4.0737
4.1131
4.1508
4.1868 **
4.2215
4.2546
4.2873
4.3183
4.3480
4.3781
4.4074
4.4376
4.4665
4.5004
4.5372
4.5787
4.6277
4.6934
4.7825
1.5106
1.4147
1.3276
1.2468
1.1719
1.1015
1.0362
0.9747
0.9169
0.8625
0.8106
0.7612
0.7134
0.6665
0.6205
0.5736
0.5254
0.4739
0.4178
0.3550
0.2809
0.1725
0
5.4127
5.3629
5.3193
5.2804
5.2456
5.2147
5.1870
5.1615
5.1385
5.1171
5.0978
5.0790
5.0614
5.0447
5.0279
5.0112
4.9919
4.9743
4.9551
4.9337
4.9086
4.8659
4.7825
* These values have been calculated in S.I. units according to DIN 1345.
** The enthalpy of liquid chlorine at 0 ◦ C was taken to be H 0 = 418.66 kJ/kg; the entropy of liquid chlorine at 0 ◦ C was taken to be
0 = 4.1868 kJ kg−1 K−1 .
The specific magnetic susceptibility at 20 ◦ C
is − 7.4 × 10−9 m3 /kg.
Liquid chlorine has a very low electrical conductivity, the value at − 70 ◦ C being
10−16 Ω−1 cm−1 . The dielectric constant of the
liquid for wavelengths greater than 10 m is 2.15
at − 60 ◦ C, 2.03 at − 20 ◦ C, 1.97 at 0 ◦ C, and
1.54 at 142 ◦ C, near the critical temperature.
Chlorine gas can be absorbed in considerable quantities onto activated charcoal and silica
gel, and this property can be used to concentrate
chlorine from gas mixtures containing it.
Chlorine is soluble in cold water, usually
less so in aqueous solutions. In salt solutions,
the solubility decreases with salt concentration
and temperature. In hydrochloric acid, chlorine
is more soluble than in water, and the solubility increases with acid concentration (Fig. 5 and
Fig. 6). In aqueous solutions, chlorine is partially
hydrolyzed, and the solubility depends on the
pH of the solution. Below 10 ◦ C chlorine forms
hydrates, which can be separated as greenish-
yellow crystals. Chlorine hydrate is a clathrate,
and there is no definite chlorine : water ratio. The
chlorine – water system has a quadruple point at
28.7 ◦ C; the phase diagram has been worked out
by Ketelaar [44].
Chlorine is readily soluble in sulfur – chlorine
compounds, which can be used as industrial solvents for chlorine. Disulfur dichloride [1002567-9], S2 Cl2 , is converted to sulfur dichloride
(SCl2 ) and sulfur tetrachloride (SCl4 ). Some
metallic chlorides and oxide chlorides, such as
vanadium oxide chloride, chromyl chloride, titanium tetrachloride, and tin(IV)chloride, are
good solvents for chlorine. Many other chlorinecontaining compounds dissolve chlorine readily. Examples are phosphoryl chloride, carbon
tetrachloride (Fig. 7), tetrachloroethane, pentachloroethane, hexachlorobutadiene (Fig. 7),
and chlorobenzene. Chlorine also dissolves in
glacial acetic acid, dimethylformamide, and nitrobenzene. The solubility of chlorine in a number of these solvents is given in Table 2.
6
Chlorine
Figure 7. Solubility of chlorine in hexachlorobutadiene
(——) and carbon tetrachloride (– – –) at 101 kPa as a
function of temperature
Figure 5. Solubility of chlorine in water, hydrochloric acid
(two concentrations), and sodium chloride solutions (three
concentrations)
All percentages are weight percents.
Figure 6. Solubility of chlorine in solutions of KCl, NaCl,
H2 SO4 , and HCl at 25 ◦ C
Table 2. Solubility of chlorine in various solvents
Solvent
Temperature,
◦
C
Solubility,
wt %
Sulfuryl chloride
Disulfur dichloride
Phosphoryl chloride
Silicon tetrachloride
Titanium tetrachloride
Benzene
Chloroform
Dimethylformamide
Acetic acid, 99.84 wt %
0
0
0
0
0
10
10
0
15
12.0
58.5
19.0
15.6
11.5
24.7
20.0
123 *
11.6 *
* g/100 cm3
3. Chemical Properties
Inorganic Compounds. Chlorine, fluorine,
bromine, and iodine constitute the halogen
group, which has marked nonmetallic properties. The valence of chlorine is determined by
the seven electrons in the outer shell. By gaining
one electron, the negatively charged chloride ion
is formed; the chloride ion has a single negative
charge and a complete shell of electrons (the argon structure). By sharing one to seven electrons
from the outer shell with other elements, the various chlorine oxidation states can be formed, for
example, in the oxides of chlorine, hypochlorites
(+ 1), chlorates (+ 5), and perchlorates (+ 7).
The bonds between chlorine and the other
halogens are mainly covalent. In the chlorine –
fluorine compounds ClF and ClF3 , there is some
ionic character to the bond, with chlorine the anion, and in the chlorine – iodine compounds ICl3
and ICl, there is some ionic character to the bond,
with chlorine the cation. Chlorine is very reactive, combining directly with most elements but
only indirectly with nitrogen, oxygen, and carbon. Excess chlorine in the presence of ammonia
salts forms the very explosive nitrogen trichloride, NCl3 . Hypochlorites react with ammonia
to produce the chloramines NH2 Cl and NHCl2 .
Chlorine
Oxygen and chlorine form several chlorine oxides (→ Chlorine Oxides and Chlorine Oxygen
Acids).
Chlorine gas does not react with hydrogen
gas [1333-74-0] at normal temperatures in the
absence of light. In sunlight or artificial light of
wavelength ca. 470 nm or at temperatures over
250 ◦ C, the two gases combine explosively to
form hydrogen chloride. The explosive limits of
mixtures of pure gases lie between ca. 8 vol %
H2 and ca. 14 vol % Cl2 (the detonation limits).
The limits depend on pressure, and the detonation range can be reduced by adding inert gases,
such as nitrogen or carbon dioxide (Fig. 8) [45,
46].
7
sulfuryl chloride, SO2 Cl2 . Under these conditions carbon monoxide and chlorine react to produce the colorless, highly toxic carbonyl chloride (phosgene), COCl2 .
Chlorine reacts with sodium cyanide and
sodium thiocyanate to produce cyanogen chloride and thiocyanogen chloride. The reaction
of chlorine with sodium thiosulfate [7772-987] (Antichlor) is used to remove free chlorine
from solutions.
Na2 S2 O3 +4 Cl2 +5 H2 O → 2 NaHSO4 +8 HCl
Chlorine reacts with carbon disulfide to produce carbon tetrachloride and disulfur dichloride.
CS2 +3 Cl2 → CCl4 +S2 Cl2
Figure 8. Explosive limits of chlorine – hydrogen – other
gas mixture
Horizontally hatched area = Explosive region with residue
gas from chlorine liquefaction (O2 , N2 , CO2 )
Checkered area = Explosive region with inert gas (N2 , CO2 )
Chlorine reacts vigorously with ammonia
3 Cl2 +4 NH3 → NCl3 +3 NH4 Cl
In the presence of the catalyst bromine, chlorine reacts with nitric oxide to give nitrosyl chloride
NO+0.5 Cl2 → NOCl
Sulfur dioxide and chlorine in the presence of
light or an activated carbon catalyst react to form
The reaction of chlorine with phosphorus
produces phosphorus trichloride (PCl3 ) and pentachloride (PCl5 ). Wet chlorine attacks most
metals to form chlorides. Although titanium
[7440-32-6] is resistant to wet chlorine, it is
rapidly attacked by dry chlorine. Tantalum is resistant to both wet and dry chlorine. Most metals are resistant to dry chlorine below 100 ◦ C,
but above a specific temperature for each metal,
combustion takes place with a flame. This specific temperature, the ignition temperature, also
depends on the particle size of the metal so that
the following values are only approximate: iron
at 140 ◦ C, nickel at 500 ◦ C, copper at 200 ◦ C,
and titanium at 20 ◦ C.
Most metal chlorides are soluble in water [3,
p. 668], notable exceptions being those of silver
(AgCl) and mercury (Hg2 Cl2 ). Chlorine liberates bromine and iodine from metallic bromides
and iodides, but is itself liberated from metal
chlorides by fluorine.
0.5 Cl2 +KBr → KCl+0.5 Br2
Selenium and tellurium react spontaneously
with liquid chlorine, whereas sulfur begins to
react only at the boiling point. Liquid chlorine
reacts vigorously with iodine, red phosphorus,
arsenic, antimony, tin, and bismuth. Potassium,
8
Chlorine
sodium, and magnesium are unaffected in liquid
chlorine at temperatures below − 80 ◦ C. Aluminum is unattacked until the temperature rises
to − 20 ◦ C, when it ignites. Gold is only slowly
attacked by liquid chlorine to form the trichloride (AuCl3 ). Cast iron, wrought iron, carbon
steel, phosphor bronze, brass, copper, zinc, and
lead are unaffected by dry liquid chlorine, even
in the presence of concentrated sulfuric acid.
Organic Compounds. The chlorine–carbon
bond is covalent in nature. Chlorine reacts with
hydrocarbons either by substitution or by addition. In saturated hydrocarbons, chlorine replaces hydrogen, either completely or partially,
to form chlorinated hydrocarbons and hydrogen
chloride. Methane can be chlorinated in stages
through to carbon tetrachloride.
In the reaction with unsaturated hydrocarbons, chlorine is added to the double or
triple bond yielding dichloro- or tetrachlorohydrocarbons, respectively. In industry the reaction velocity is increased by light (photochlorination), heat (cracker furnace), or catalysts. In
aromatic hydrocarbons, both addition and substitution are possible, depending on conditions
(light, temperature, pressure, or catalysts).
Chlorine Compounds in Nature. More
than 1800 chlorine-containing compounds have
been identified in nature at present. These organic chlorine compounds, which range from
chloromethane to the complex antibiotic vancomycin, are produced by marine and terrestrial
plants, bacteria, fungi, lichen, insects, marine
animals, some higher animals, and a few mammals. For example, the annual natural production of chloromethane is estimated to exceed
4 × 106 t/a [47].
4. Chlor-Alkali Process
In the chlor-alkali electrolysis process, an aqueous solution of sodium chloride is decomposed electrolytically by direct current, producing chlorine, hydrogen, and sodium hydroxide
solution. The overall reaction of the process
2 NaCl+2 H2 O → Cl2 +H2 +2 NaOH
Figure 9. Flow diagram of the chlor-alkali mercury process
takes place in two parts, at the anode and at the
cathode. The evolution of chlorine takes place at
the anode:
2 Cl− → 2 Cl+2 e− → Cl2 +2 e−
There are three basic processes for the electrolytic production of chlorine, the nature of
the cathode reaction depending on the specific
Chlorine
process. These three processes are (1) the diaphragm cell process (Griesheim cell, 1885),
(2) the mercury cell process (Castner – Kellner
cell, 1892), and (3) the membrane cell process
(1970).
Each process represents a different method of
keeping the chlorine produced at the anode separate from the caustic soda and hydrogen produced, directly or indirectly, at the cathode.
9
These three processes are described in detail
in the following three chapters. The basic flow
sheets of the three processes are shown in Figures 9, [10]], [11]]. In all three processes, nearly
saturated, purified brine is introduced into the
electrolysis cell.
Figure 10. Flow diagram of the chlor-alkali diaphragm
process
Figure 11. Flow diagram of the chlor-alkali membrane process
(* optional)
10
Chlorine
Figure 12. Processing of hydrogen gas from the amalgam decomposer
a) Vertical decomposer; b) Individual cell hydrogen cooler; c) Safety seal; d) Demister; e) Blower; f) Final hydrogen cooler;
g) Mercury removal equipment
The hydrogen produced is cooled as it leaves
the decomposer or the cathode compartment
and is carried through electrically insulated
pipework to a vessel fitted with a water seal
(Fig. 12). If a hydrogen – air mixture forms because of a shutdown or breakdown, the seal
allows the mixture to escape. A demister ensures that the gas is free of spray, whether water or sodium hydroxide solution. The hydrogen
is compressed by Roots-type blowers or reciprocating compressors before it passes through
coolers on its way to the consuming plants. At
no stage is the pressure allowed to fall below
ambient pressure.
Electrolytic hydrogen is very pure, > 99.9
%; however, unwanted traces of oxygen can be
removed by reaction with the hydrogen over a
platinum catalyst. The hydrogen is used for organic hydrogenation, catalytic reductions, ammonia synthesis and to provide hot flames or protective atmospheres in welding technology, metallurgy, or glass manufacture. It is also used in
the manufacture of high-purity hydrogen chloride by combustion with chlorine and as a fuel
for heating and drying.
In the mercury cell process, sodium amalgam
is produced at the cathode. The amalgam is reacted with water in a separate reactor, called the
decomposer, to produce hydrogen gas and caustic soda solution.
Because the brine is recirculated, solid salt
is required for resaturation. The brine, which
must be quite pure, is first dechlorinated and
then purified by a straightforward precipitation –
filtration process.
The products are extremely pure. The chlorine, along with a little oxygen, generally can be
used without further purification. The sodium
hydroxide solution contains little chloride and
leaves the decomposer with a 50 wt % concentration.
Of the three processes, the mercury process
uses the most electric energy; however, no steam
is required to concentrate the caustic solution.
The use of large quantities of mercury demands
measures to prevent environmental contamination. In addition, the hydrogen gas and sodium
solution must be freed from mercury. Generally,
the operation of the cells is not simple.
In the diaphragm cell process, the anode area
is separated from the cathode area by a permeable, generally asbestos-based diaphragm. The
brine is introduced into the anode compartment and flows through the diaphragm into the
cathode compartment. Cheaper solution-mined
brine can be used; the brine is purified by precipitation – filtration.
A caustic brine leaves the cell, and this brine
must be freed from salt in an elaborate evaporative process. Even so, the resultant 50 wt %
sodium hydroxide solution contains up to 1 wt %
NaCl. The salt separated from the caustic brine
can be used to saturate dilute brine. The chlorine contains oxygen and must be purified by
liquefaction and evaporation.
The consumption of electric energy with the
diaphragm cell process is ca. 15 % lower than
for the mercury process, but the total energy consumption is higher because of the steam required
to concentrate the caustic brine (see Fig. 62). Environmental contamination with asbestos must
Chlorine
be avoided. Under constant operating conditions, cell operation is relatively simple.
In the membrane cell process, the anode and
cathode are separated by a cation-permeable ionexchange membrane. Only sodium ions and a
little water pass through the membrane.
As in the mercury process, the brine is dechlorinated and recirculated, which requires solid
salt to resaturate the brine. The life of the expensive membrane depends on the purity of the
brine. Therefore, after purification by precipitation – filtration, the brine is also purified with an
ion exchanger.
The caustic solution leaves the cell with a
concentration of 30 – 36 wt % and must be concentrated. The chloride content of the sodium
hydroxide solution is almost as low as that from
the mercury process. The chlorine gas contains
some oxygen and must be purified by liquefaction and evaporation.
The consumption of electric energy with the
membrane cell process is the lowest of the three
processes, ca. 25 % less than for the mercury
process, and the amount of steam needed for
concentration of the caustic is relatively small
(see Fig. 62). The energy consumption should
be even lower when oxygen-consuming electrodes become common. There are no special
environmental problems. The cells are easy to
operate and are relatively insensitive to current
changes, allowing greater use of the cheaper offpeak-time electric power.
4.1. Brine Supply
The brine used in the mercury cell and membrane cell processes is normally saturated with
solid salt although there are some installations
that use solution-mined brine on a once-through
basis. The brine supply for diaphragm cells is
always used on a once-through basis, although
the salt recovered from caustic soda evaporators
may be recycled into the brine supply.
Salt. The basic raw material for the mercury
cell and membrane cell processes is usually solid
salt. This may be obtained from three sources:
rock salt, solar salt, or vacuum-evaporated salt
from purifying and evaporating solution-mined
brine.
11
In the United States and Europe, rock salt is
most commonly used. The most important impurities are shown in Table 3. The concentrations
of these impurities depend on the method of production and on the different grades: crude rock
salt, prepared rock salt, and evaporated salt. Solar salt is used in Japan and many other parts
of the world, the most important sources being
Australia, Mexico, China, Chile, India, and Pakistan. The salt produced by solar evaporation is
usually much less pure than rock salt. In a few
cases the salt may be obtained from other processes, such as caustic soda evaporation in the
diaphragm process.
A new upgrading process (Salex) has been
developed by Krebs Swiss [48]. It removes the
impurities by selective cracking of the salt crystals and a washing process. Salt losses are minimized, and the purity exceeds 99.95 % NaCl.
Table 3. Impurities in rock salt and sea salt, wt %
Insolubles
Water
Calcium
Magnesium
Sulfate
Potassium
Rock salt
Sea salt
≤2
≤3
0.2 – 0.3
0.03 – 0.1
≤0.8
≤0.04
0.1 – 0.3
2.0 – 6.0
0.1 – 0.3
0.08 – 0.3
0.3 – 1.2
0.02 – 0.12
Brine Resaturation. In older plants, the
open vessels or pits used for storing the salt
are also used as resaturators. The depleted brine
from the cells is sprayed onto the salt and is saturated, the NaCl concentration reaching 310 –
315 g/L. Modern resaturators are closed vessels,
to reduce environmental pollution [49], which
could otherwise occur by the emission of a salt
spray or mist. The weak brine is fed in at the
base of the resaturator, and the saturated brine is
drawn off at the top. If the flow rates of the brine
and the continously added salt are chosen carefully, the differing dissolution rates of NaCl and
CaSO4 result in little calcium sulfate dissolving within the saturator [50]. Organic additives
also reduce the dissolution rate of calcium sulfate [51]. The solubility (g per 100 g of H2 O) of
NaCl in water does not increase much with temperature (t, ◦ C), whereas the solubility of KCl
does:
12
Chlorine
t
0
20
40
60
80
100
cNaCl
cKCl
35.6
28.2
35.8
34.4
36.4
40.3
37.0
45.6
38.5
51.0
39.2
56.2
Brine Purification. In mercury cells, traces
of heavy metals in the brine give rise to dangerous operating conditions (see page 15), as does
the presence of magnesium and to a lesser extent calcium [52]. In membrane cells, divalent
ions such as Ca2+ or Mg2+ are harmful to the
membrane. The circulating brine must be rigorously purified to avoid any buildup of these substances to undesirable levels [7]. Calcium is usually precipitated as the carbonate with sodium
carbonate; magnesium and iron, as hydroxides
with sodium hydroxide; and sulfate, as barium
sulfate.
The reagents are usually mixed with weak
brine and added to the brine stream at a controlled rate. If solar salt is used, treatment costs
may be reduced by prewashing the salt [53]. In
order to precipitate calcium at low pH, sodium
bicarbonate [54] or phosphoric acid [55] can be
added.
The sulfate content can be reduced without
the use of expensive barium salts by discharging a part of (purging) the brine [56], by crystallization of Na2 SO4 · 10 H2 O on cooling the
brine stream [57], by precipitation of the double
salt Na2 SO4 · CaSO4 [58], by an ion-exchange
process, or by membrane nanofiltration [59].
Hoechst [60] has a process for recovering barium sulfate of pure pigment quality by precipitation under acid conditions. Chlorate buildup can
be avoided by addition of sodium metabisulfite
Na2 S2 O5 [61].
After stirring for 1 – 2 h, the precipitated impurities are removed by filtration alone or by
sedimentation followed by filtration. Sedimentation is carried out in large circular settling tanks,
from which the slurry is removed by mechanical
raking equipment, e.g., Clariflocculator, Cyclator, or Dorr thickener. Filtration is carried out
with a sand filter, a pressure-leaf filter with filter cloths of chlorine-resistant fabrics, or candle filters automatically cleaned by backflow of
brine. The filter is cleaned by water jets, vibrating, or shaking. The separated filter cake is concentrated to 60 – 80 % solids content in rotary
drum vacuum filters or centrifuges before disposal. Any soluble material present may be re-
moved from the sludge by washing with water.
Barium salts may be recovered by treating with
sodium carbonate under pressure [62]. The purified brine should contain ideally cCa < 2 mg/L,
cMg < 1 mg/L, and cSO4 < 5 g/L.
In the diaphragm process, the removal of sulfate is not always necessary because SO2−
4 can
be removed from the cell liquor as pure Na2 SO4
during the concentration process. In the membrane process, the brine must be purified to a
much higher degree to avoid the deterioration
of the membrane. The Ca2+ and Mg2+ concentration must be < 0.02 ppm (20 ppb), so a second, fine purification step is required (see Section 7.2.1).
Before the brine enters the electrolysis cells,
it should be acidified with hydrochloric acid to
pH < 6, which increases the life of the titanium
anode coating, gives a purer chlorine product
with higher yield, and reduces the formation of
hypochlorite and chlorate in the brine.
Brine Dechlorination. In the mercury and
membrane processes, the depleted brine leaving the cells must be dechlorinated before resaturation. Further acidification with hydrochloric
acid to pH 2 – 2.5 reduces the solubility of chlorine by shifting the equilibrium point of hydrolysis and inhibits the formation of hypochlorite
and chlorate. Chlorine discharged in the anolyte
tank prior to dechlorination may be fed into the
chlorine system. The dissolved chlorine of the
brine then is still 400 – 1000 mg/L, depending on
pH and temperature. The brine is passed down a
packed column or sprayed into a vacuum of 50
– 60 kPa, which reduces the chlorine concentration in the brine to 10 – 30 mg/L. The vacuum
is produced by steam jet or liquid-ring vacuum
pump. The pure chlorine gas obtained is fed into
the chlorine stream.
The water that evaporates from the dechlorinated brine is condensed in a cooler. The condensate, which may be chemically dechlorinated, is
returned to the brine circulation system if necessary to maintain the volume of the brine circuit.
If necessary, the remaining chlorine content can
be further reduced by blowing with compressed
air, by a second vacuum treatment, by treatment
with activated carbon [63], or by chemical treatment with hydrogen sulfite, thiosulfate, sulfur
dioxide, or sodium hydrogensulfide.
Chlorine
Brine Monitoring. The sodium chloride
concentration in the brine is determined by density measured by equipment involving radioactive isotopes, vibration techniques, hydrometry,
or weighing.
The pH following alkali or acid additions is
determined with glass electrodes, and the redox
potential following chlorine removal is determined with metal electrodes. Excess OH− and
CO2−
3 ions ensure adequate precipitation of dissolved calcium, iron, and magnesium. After filtration, a test sample of 100 mL should require
4 – 6 mL of 0.1 N acid to reach the phenolphthalein end point and a further 0.5 – 1.5 mL to
reach the methyl orange end point. Inadequate
filtration is detected by turbidimetry in transmitted light or by the Tyndall effect. Calcium and
magnesium are determined hourly, and chlorate
and sulfate about once per day, all by titration.
13
Each set of rectifiers is connected through
high-voltage switchgear to the three-phase supply [65]. Smaller units use a 10 – 30 kV supply,
but large units can be connected into the highvoltage power system (> 100 kV) [66].
The unit cost of the d.c. supply decreases
with increasing voltage and current. A plant is
therefore most economical when as many highcurrent cells as possible are connected in series
[67]. Total currents of 450 000 A are achieved.
The switches for short-circuiting the cells are designed for 10 000 – 30 000 A and are operated by
compressed air, hydraulically, or by spring action. Erosion of the main contacts is dealt with
by using replaceable pre-contacts [68]. The contacts are protected from corrosion by installation
in vacuum housings.
The current in the bus bars or in anode rods
can be measured by means of iron-free transportable equipment with an accuracy of ca. 1 %
[69].
4.2. Electricity Supply
Since 1960 the direct current for electrolysis has
been provided exclusively by silicon rectifiers.
A set of rectifiers can supply up to 450 000 A.
Voltages up to 4.0 kV per diode are feasible, but
usually for safety, a peak a.c. voltage of 1500 V,
corresponding to a d.c. output of 1200 V, is not
exceeded. Liquid cooling of the diodes permits
a compact design, and self-contained equipment
reduces leakage losses. Modern membrane cell
plants also use continuously variable thyristor
converters in place of silicon diodes [64].
Rectification equipment is required to provide steady direct current at a voltage determined
by the cell room. The current must remain steady
even though the voltage is varied both by the operating condition of the cells and by the number
of cells operating. The rectifier equipment usually consists of
transformer capable of variable output voltage with adequate compensation for changing
input voltage
silicon rectifiers or thyristors
constant-current control gear
transducers for metering and control
control panels
isolators
cooling equipment
ancillary safety and monitoring equipment
5. Mercury Cell Process
The clean separation of chlorine from the cathode products is possible because of the high
overvoltage of hydrogen at the mercury electrode. Hydrogen and sodium hydroxide are not
produced at the cathode; instead, sodium is produced and dissolves in the mercury as an amalgam. The liquid amalgam is removed from the
electrolytic cell to a separate reactor, called the
decomposer or denuder, where it reacts with
water in the presence of a catalyst to form the
sodium hydroxide and hydrogen gas. The process may also be used to produce potassium hydroxide by feeding the cell with potassium chloride solution, although this is much less common. The sodium hydroxide is produced from
the denuder at a concentration of ca. 50 wt %;
the maximum value is 73 wt %. The hydroxide
solution is very pure and almost free from chloride contamination.
The process was developed in 1892 almost simultaneously by H. Y. Castner and C. Kellner
and used on an industrial scale, although the
amount of chlorine produced remained relatively small until 1930, when the rapid growth
of the rayon (artificial silk) industry, especially
in Germany, increased the demand for pure
14
Chlorine
chloride-free sodium hydroxide solution. At this
time, the horizontal high-current cell was developed and output increased rapidly. The development work in Germany was described in the
FIAT final reports, published after World War II,
and this led to widespread use of the process in
Europe and Japan [70]. In the United States, the
mercury cell process became more widespread,
increasing its share of chlorine production from
3 – 4 % in 1945 to 20 % in 1960, reaching a maximum of 27 % in 1970.
The development of the mercury cell can be
followed in the technical data: the cell current
increased from 3.4 kA in 1895 to ca. 30 kA in
1945, 200 kA in 1960, and 450 kA in 1970. The
current density rose from 2 kA/m2 in 1950 to the
current maximum of 15 kA/m2 . The cell area increased over the same period from ca. 7 m2 to
37.5 m2 , while the k-factor ( specific voltage coefficient, see page 16) was reduced by 50 %.
Since 1972 the importance of the mercury
cell has decreased. Increasing concern about the
effect of mercury on the environment has led to
a considerable increase in the number and variation of statutory regulations that affect the mercury cell process.
In particular, widespread concern about cases
of mercury poisoning in Japan, which were not
related to the mercury cell process, caused the
process to be legally banned since 1972. However, conversion to the alternative processes was
delayed because of demand for low-chloride
sodium hydroxide and because of the anticipated
advantages of the rapidly developing membrane
cell process. The last remaining mercury cell installations for NaCl were closed in 1986.
In the other countries, existing mercury cell
plants are still in operation, but official regulations and uncertainty about possible further legal
restrictions have hindered expansion.
In Europe and the United States great efforts are being made to develop methods of protecting the environment from mercury (see Section 5.3.5). These measures have greatly reduced
emissions of mercury into the atmosphere and
into wastewater to the extent that the present levels of emitted mercury are negligible in comparison to those arising from natural sources, such
as volcanic action, geological erosion, or other
nonnatural sources such as fuel combustion or
metallurgical processes.
In 1984, the mercury cell process accounted
for 45 % of world chlorine production [71].
Since then no new plants have been built. In
the coming decades most of the existing mercury cell plants will be shut down or converted to membrane cell technology. Only plants
with speciality products such as extremely pure
sodium hydroxide, potassium hydroxide, alcoholates, and dithionites will use the mercury process in future. These plants will meet the highest
emission control standards.
5.1. Principles
The cathode reaction
Na+ +e− +Hgx → NaHgx
forming sodium amalgam, is followed by the
decomposition reaction in a separate reactor
2 NaHgx +2 H2 O → 2 NaOH+H2 (g) +2 Hgx
Process Description
(Fig. 13). Mercury
flows down the inclined base of the electrolytic
cell (A). The base of the cell is electrically connected to the negative pole of the d.c. supply. On
top of the mercury and flowing cocurrently with
it is a concentrated brine with a sodium chloride content of ca. 310 g/L at the inlet. Anodes
are placed in the brine so that there is a small
gap between the anode and the mercury cathode.
The concentration of the amalgam is maintained
at 0.2 – 0.4 wt % Na, so that the amalgam flows
freely (Fig. 14). The chlorine gas and depleted
brine (270 g/L) flow out of the cell, either separately or as a two-phase mixture separated later
in the process. The amalgam flows out of the
cell through a weir and into the decomposer. The
amalgam may be passed through a water wash
between the cell and the decomposer to remove
traces of sodium chloride. The amalgam flows
through the decomposer countercurrent to a flow
of softened or demineralized water in the presence of a catalyst to produce sodium hydroxide
solution and hydrogen. Stripped of its sodium,
the mercury flows out of the lower end of the
decomposer and is recirculated through a pump
back into the cell.
Chlorine
15
Figure 13. Schematic view of a mercury cell with decomposers
A) Mercury cell: a) Mercury inlet box; b) Anodes; c) End box; d) Wash box
B) Horizontal decomposer: e) Hydrogen gas cooler; f) Graphite blades; g) Mercury pump
C) Vertical decomposer: e) Hydrogen gas cooler; g) Mercury pump; h) Mercury distributor; i) Packing pressing springs
Figure 14. Freezing point curves of sodium amalgam and
potassium amalgam
Anode Reactions. The oxidation of chloride
ions to chlorine gas has a standard potential of
1.358 V. In a 300 g/L sodium chloride solution
at 70 ◦ C, the reversible reaction potential is reduced to 1.248 V [15, p. 339]. Some side reactions occur, such as the oxidation of OH− and
SO2−
ions and the electrochemical formation
4
of chlorate ions. Nonelectrochemical reactions
also take place in the region of the anode, such
as hypochlorite formation (because of hydrolysis of chlorine) and chlorate formation. All of
these side reactions represent a loss of efficiency.
Cathode Reactions. The standard potential
of the hydrogen-liberating reaction is 0 V, which
is considerably higher than the potential for
the formation of 0.2 wt % sodium amalgam,
− 1.868 V. However, hydrogen is not liberated
at the mercury surface because the reaction is
kinetically inhibited. Mainly sodium ions are
discharged. At the sodium chloride concentrations used, the reversible potential is reduced by
ca. 0.2 V. (Exact values of the discharge potential are given as a function of the sodium concentration in the amalgam, the sodium chloride
concentration in the brine, and the temperature
[72].) Electrochemical side reactions occur: the
reduction of chlorine molecules or hypochlorous acid and the liberation of hydrogen gas. In
addition, sodium in the amalgam can react directly with free chlorine, or chlorite and chlorate
ions can be reduced to chloride by the action of
nascent hydrogen at the cathode. All of these
side reactions represent a loss of efficiency, normally ca. 2 – 4 % under good operating conditions.
Contamination of the system by heavy metals
can lead to a reduction of the hydrogen discharge
potential at the mercury cathode, thus increasing
hydrogen liberation, and reducing amalgam formation [73]. The hydrogen concentration in the
chlorine can increase to the point at which the
cell and downstream chlorine handling equipment contains explosive mixtures. The probability of such problems is estimated by a hazard
analysis of an existing plant [74 – 76].
16
Chlorine
The cell system is sensitive to trace quantities of catalysts in the brine, for example, vanadium, molybdenum, and chromium at the 0.01 –
0.1 ppm level or iron, cobalt, nickel, and tungsten at the parts per million level. Magnesium,
calcium, aluminum, and barium are also active
at the parts per million level.
In addition, relatively high concentrations of
sodium in the amalgam (> 0.5 wt %) can cause
increased hydrogen evolution in the cells. Potassium chloride electrolysis is considerably more
sensitive to both catalysts and high concentration in the amalgam than the sodium chloride
process.
Current Efficiency. The theoretical electrochemical equivalents representing the materials produced or consumed in the electrolysis of
sodium chloride or potassium chloride brines
are given in Table 4. In practice, the yield is
ca. 95 – 97 % of the theoretical value, owing to
side reactions at the electrodes and in the electrolyte. With activated titanium anodes, the yield
is largely independent of the distance between
the electrodes.
The decrease in salt concentration ∆c is determined by the current I, the brine flow rate M,
and the electrochemical equivalent f.
∆c = f I/M
The usual units are c in g/L, f in kg kA−1 h−1 ,
I in kA, and M in m3 /h.
Cell Voltage. The d.c. voltage across the cell
circuit is determined by five factors:
1) The reversible decomposition voltage of the
salt
2) The overpotentials of the chlorine and alkali
metal at the electrodes
3) The voltage drop in the electrolyte
4) Voltage losses in the bus bars, switches, electrical conductors, anode materials, and cathode
5) The operating current density of the cells
Factor 1.The reversible decomposition potential of NaCl under standard conditions
is E ◦ = 3.226 V (KCl E ◦ = 3.234 V). Under
the operating conditions cNaCl = 290 g/L,
camalgam = 0.15 %, and 70 ◦ C, the reversible decomposition voltage is E = 3.095 V [77].
Factor 2. The overpotential of chlorine depends on the material and shape of the anodes. At
the high current densities (10 kA/m2 ) present in
modern cell rooms, the overpotential can reach
several hundred millivolts, outweighing the effect of concentration changes in the electrolyte
(concentration polarization) [78] and the retarding effect that formation of molecular chlorine
has on the process of ion discharging [79]. Chlorine gas bubbles cover part of the anode surface,
thereby increasing the current density at the free
surface. The anode is designed so that the gas
bubbles are liberated as quickly as possible. The
rapid removal of these gas bubbles from the reaction zone is one of the advantages of titanium
anodes over graphite anodes [80].
The overpotential of sodium on the amalgam
cathode is caused by the limited diffusion rate
of the liberated sodium atoms into the amalgam,
but it is small compared to the chlorine overvoltage.
Factor 3. The specific conductivity of
sodium chloride solutions increases with concentration and temperature (Fig. 15), but is independent of pH over the range 2 – 11. The brine
normally enters the cells at 60 – 70 ◦ C and leaves
the cells at 75 – 85 ◦ C. The conductivity of potassium chloride solutions at 70 ◦ C is 30 % greater
than that of sodium chloride solutions. Chlorine
gas bubbles in the electrolyte increase the resistance between anodes and cathode. Better circulation of the electrolyte in the gap between electrodes allows more rapid removal of gas bubbles,
thus reducing the voltage.
Factor 4.The voltage losses in the cell room
are minimized by compactly arranging the cells,
which shortens the current path. The relatively
low conductivity of steel cell bases can be improved by copper or aluminum fittings. These
measures also reduce problems caused by magnetic fields, which occur in wide cells [81].
Factor 5. In practice, cell current density and
cell voltage have a linear relationship. The slope
of the line is termed the specific voltage coefficient or k-factor, a useful measure of the specific
energy requirement of cells produced by different manufacturers.
Chlorine
17
Table 4. Electrochemical equivalents f, kgkA−1 h−1
Element
Element produced
Salt required
Alkali produced
Na
K
Cl2
H2
0.8580
1.4586
1.3228 (0.4115 m3 STP)
0.0376 (0.4185 m3 STP)
2.1810 (NaCl)
2.7816 (KCl)
1.4923 (NaOH)
2.0931 (KOH)
In addition to the d.c. voltages considered
above, there are energy losses across the transformer and rectification equipment. All cell installations use a.c. power, which is rectified by
silicon diodes in which the energy losses are
minimized by operating at greater than 100 V.
This voltage is achieved by operating at least 25
cells in series.
Energy Consumption. To operate a cell installation economically, the consumption of d.c.
electrical energy per unit mass of product must
be minimized. The specific energy consumption
w is given by
w = 1000 Ucell /a f
Figure 15. Specific conductivity of sodium chloride
solutions
The cell voltage is given by U cell = 3.15 + kJ,
J = current density, kA/m2 , k = specific voltage
coefficient, V m2 kA−1 .
Computer-controlled cells with activated titanium anodes are run with k-factors from 0.085
to 0.11. The corresponding cell voltages at
10 kA/m2 are 4.00 – 4.25 V (Fig. 16).
Figure 16. Cell voltage and specific energy consumption
per tonne of Cl2 versus cell current density
where w = kWh/t, a = yield factor or current efficiency, f = electrochemical equivalent,
kg kA−1 h−1
For example, if the cell voltage U cell is 4.20 V
and the current efficiency is 0.970, then ca.
3275 kWh is required to produce 1 t of chlorine + 1.13 t of NaOH. Since af is almost a constant, the specific energy consumption per tonne
of chlorine w is effectively proportional to the
cell voltage. In that case, w also depends on the
cell current density (see Fig. 16). In the example,
w = 3275 kWh corresponds to 10 kA/m2 .
The total energy requirement per tonne of Cl2
must also include the transformer and rectifier
losses (30 – 40 kWh/t) and the energy requirements of all of the ancillary equipment (120 –
160 kWh/t).
A mathematical model of the cell has been
described [82].
Decomposition of the Amalgam. The
amalgam is decomposed in horizontal decomposers, alongside or beneath the cell, or more
often since ca. 1960 in vertical decomposers
or denuders. The energy stored in the amalgam
18
Chlorine
has an emf of ca. 0.8 V. The hydrogen overpotential at the amalgam prevents spontaneous
decomposition in contact with water, and a catalyst (depolarizer) must be used. The overall
decomposition reaction is
2 NaHgx +2 H2 O → 2 NaOH+H2 +2 Hgx
and takes place in two stages, first as an anode
reaction at the surface of the amalgam
2 Na → 2 Na+ +2 e−
and then as a cathode reaction on the catalyst
surface, where the water is decomposed
2 H2 O+2 e− → 2 OH− +H2
Industrial decomposers are essentially
short-circuited electrochemical primary cells
(Fig. 17). The most common catalyst is graphite
[7782-42-5], usually activated by oxides of iron,
nickel, or cobalt or by carbides of molybdenum or tungsten. The hydrogen overpotential
on graphite (0.5 – 0.6 V at 2 kA/m2 and 80 ◦ C)
increases with current density and decreases
with temperature; therefore, the decomposer
should be operated at as high a temperature
as possible [83]. Good catalyst material must
meet many requirements: resistance to alkali
solutions, hydrogen, and mercury; low hydrogen overpotential; good electrical conductivity;
long-lasting activity; wettability by amalgam;
and incapability of amalgamation.
Attempts to recover some of the energy stored
in the amalgam by creating an electrical circuit
by using the catalyst as the anode separated from
the amalgam or by using the amalgam electrode
with an oxygen gas diffusion electrode have so
far had no practical outcome [84].
Horizontal decomposers are ducts with a
rectangular cross section, which are installed
with a 1 – 2.5 % slope near to or underneath the
cells. The amalgam flows in a stream ca. 10 mm
in depth, and the sodium content is thereby reduced to < 0.02 wt %. The catalyst consists of
graphite blades 4 – 6 mm thick, which are immersed in the amalgam in a lengthwise direction
(Fig. 18, also see Fig. 13 B). The water for the
reaction, which is softened or demineralized by
ion exchange, flows in the direction opposite the
amalgam and is removed as 50 % caustic alkali
solution. The hydrogen gas is cooled as it leaves
the decomposer so that any condensed water and
mercury run back into the decomposer. Advantages of the horizontal decomposer are serviceability, simple construction, and a pure product
that is low in mercury. However, horizontal decomposers require a greater mercury inventory
than vertical decomposers.
Figure 18. Cross section through a horizontal decomposer
a) Amalgam; b) Bolt; c) Graphite blades; d) Hydrogen;
e) Sodium hydroxide solution; f) Decomposer casing;
g) Spacers
Vertical decomposers are designed as towers
[85] containing packings of activated graphite
spheres or other shapes 8 – 20 mm in diameter. The towers are packed 0.6 – 0.8 m high. The
cross section of the tower is 0.35 m2 per 100 kA
of cell current. The amalgam is fed in via an
overhead distributor, and the mercury is pumped
Figure 17. Principle of amalgam decomposition
Chlorine
from the base of the tower back to the cell by
a closed centrifugal pump (see Fig. 13 C). The
water for the reaction is fed into the base of the
tower and flows upward counter to the amalgam. The 50 % caustic alkali solution flows out
at the top. The smaller volume of the vertical
decomposer leads to higher product temperature because of the greater energy intensity of
the system. Cooling the hydrogen is essential.
Compared with the horizontal decomposer, the
amount of space required is small, and the mercury inventory is small, but the caustic alkali
contains more mercury.
In alternative decomposition reactions, other
products may be obtained from the amalgam
in place of sodium or potassium hydroxide solutions [10, p. 518], [86]: sodium sulfide from
sodium polysulfide solution, alcoholates from
alcohols, sodium dithionite from sodium hydrogen sulfite, hydrazobenzene or aniline from nitrobenzene, adiponitrile from acrylonitrile, and
alkali metals by distillation.
19
rubberized fabric. The anodes, generally made
of activated titanium, hang in groups from carrying devices that can be varied in height manually, hydraulically, or by motor-operated lifting
devices. Each cell can be short-circuited externally by a switch. The cell bus bars are usually
copper. The anodes are protected from internal
short circuits by means of electronic monitoring systems. The size of the cells can be varied
within a broad range to give the desired chlorine
production rate. Computer programs optimize
the cell size, number of cells, and optimum current density as a function of the electricity cost
[87] and capital cost.
For comparison, a list of cells manufactured
by leading engineering firms and cell characteristics is given in Table 5. Cathode surface areas
are ca. 17 – 30 m2 , and nominal currents are ca.
170 – 300 kA [32, p. 204].
5.2.1. Uhde Cell
5.2. Mercury Cells
During the first decades after the rocking cells of
Castner and Kellner were first commissioned,
considerable efforts were made to develop suitable materials for the cells and the anodes. A
large number of cell configurations were tested,
resulting in the development of the continuous
cell. Since 1950, the cell areas and the specific
load were increased considerably.
In 1972, the changeover from graphite to
metallic anodes began, with a parallel development of computer monitoring and control, leading to improved short-circuit protection and a
reduction of the specific energy consumption by
computer-controlled anode adjustment, of great
significance in view of the drastic increase in
electricity costs in the late 1970s. In the years
following 1972, producers operating the electrolysis plants also concentrated on the development and installation of devices to reduce mercury emissions.
The cells currently available possess a number of common features. The mercury flows over
a steel base that has a slope of 1.0 – 2.5 %. The
flanged side walls are lined with rubber. The cell
covers are mostly steel, lined with rubber or titanium on the underside, but they may also be
The Uhde cells (Fig. 19, also see Fig. 22) are
available with a cathode surface area 4 – 30 m2
for chlorine production rates 10 – 1000 t/d for the
complete cell installation. The brine flows in via
an inlet box fitted with two pipes for the removal
of chlorine. The weak brine is removed at the
end of the cell. The solid cover is fixed to the
side walls by clamps. The anodes are suspended
in groups in carrying frames supported near the
cells on transverse girders with lifting gear. The
anode rods are raised and lowered within a bellows seal. Short copper bus bars between the
cells also serve for shunt measurement of the anode currents. The electric current is brought in
above the cell covers via flexible copper straps
that run immediately above the anode rods and
are bolted to them. The compressed-air switches
are situated under the cells. The cell bottom is
usually a current conductor when cells are shortcircuited, but in wide, heavily loaded cells the
cathode current is carried by copper bus bars to
prevent the occurrence of strong magnetic fields,
which could interfere with the amalgam flow.
The automatic equipment for protection and adjustment of the anodes depends on the shunt
measurement of the currents and is controlled
by a central computer. In this way, an optimum
k-factor is selected for each cell. The vertical
20
Chlorine
Table 5. Characteristics of modern mercury cells
Characteristic
Cell type
Cathode area, m2
Cathode dimensions, l×b, m2
Slope of cell base, %
Rated current, kA
Max. current density, kA/m2
Cell voltage at 10 kA/m2 , V
Number of anodes
Stems per anode
Number of intercell bus bars
Quantity of mercury per cell, kg
Energy requirement per tonne of Cl2 , kWh d.c.
Manufacturer
Uhde
De Nora
Olin – Mathiesen Solvay
Krebs Paris
300 – 100
30.74
14.6×2.1
1.5
350
12.5
4.25
54
4
36
5000
3300
24M2
26.4
12.6×2.1
2.0
270
13
3.95
48
4
32
4550
3080
E 812
28.8
14.8×1.94
1.5
288
10
4.24
96
2
24
3800
3300
15 KFM
15.4
9.6×1.6
MAT 17
17
12.6×1.8
1.7
170
10
4.10
96
1
24
3200
160
10.4
4.30
24
4
12
1650
3400
Figure 19. Uhde mercury cell
a) Cell base; b) Anode; c) Cover seal; d) Cell cover; e) Group adjusting gear; f) Intercell bus bar; g) Short-circuit switch;
h) Hydrogen cooler; i) Vertical decomposer; j) Mercury pump; k) Anode adjusting gear; l) Inlet box; m) End box
decomposers are provided with hydrogen coolers and are situated at the end of each cell. The
amalgam flows into the decomposer under the
force of gravity [88].
5.2.2. De Nora Cell
The size of the De Nora cell (Fig. 20, also see
Fig. 23) varies from 4.5 to 36 m2 , corresponding
to electric currents from 45 to 400 kA. The cover
is a flexible multilayer sheet of elastomer spread
over the cell trough. This cover is supported by
the anode rods and seals them. The DSA anodes
(see Section 8.1) are held rigidly in strong carrying frames, which are automatically adjusted
by electric motors. Individual anode adjustment
is not provided. The anode rods are individually
connected by flexible copper straps to the an-
Chlorine
21
Figure 20. Cross section through the De Nora mercury cell
a) Cell base (steel); b) Side wall (rubber-lined steel); c) Lifting gear; d) Transverse support; e) Lengthwise support; f) Anode
carrier; g) Anode rod; h) Anode surface; i) Adjusting motor; k) Bus bar; l) Flexible anode current strap; m) Multilayer cell
cover; n) Service walkway; o) Intercell bus bar; p) Switch; q) Insulator; r) Switch drive; s) Support
ode bus bars. The cathode current is carried by
copper bus bars. Devices for the improvement
of brine circulation and gas removal within the
cells reduce specific energy consumption. Consequently, the reduction in brine concentration
can be increased from the usual 35 – 40 g/L to
60 – 70 g/L, and the brine circulating rate can be
reduced by ca. 40 %. Separate outlets are present
at the inlet box for the normal chlorine gas production and the weak chlorine gas produced during start-up. The graphite catalyst in the vertical
decomposer is activated with molybdenum.
5.2.3. Olin – Mathieson Cell
The special feature of the Olin – Mathieson cell
lies in the system of mounting and adjusting the
anodes. Above each row of anode rods, a Ushaped copper or aluminum bus bar also serves
to support the anode lifting gear. The anode rods
are bolted to the U-shaped bus bar. The anodes
are adjusted as a group, either manually or by a
remote computer with the remote computerized
anode adjuster (RCAA) system. The currents are
measured independently of the cell potentials by
means of reed contacts [89].
5.2.4. Solvay Cell
The bus bars in the Solvay cells are made primarily of aluminum. Above the cells is a cover that
also serves as a convenient walkway, giving access to the anode rods. The titanium anodes are
specially coated and are automatically adjusted
by computer. The tall vertical decomposers are
located under the cells.
5.3. Operation
The aspects of the operation of mercury cells that
typically differ from those of the other processes
are the brine circulation system, the cell room,
treatment of the products, measurement and control, and reduction of mercury emissions.
5.3.1. Brine System
A typical brine circulation system for the mercury cell process is shown in Figure 21. In the
cells the sodium chloride concentration of the
brine is reduced by 35 – 60 g/L to 260 – 280 g/L
22
Chlorine
Figure 21. Schematic diagram of a brine circulation system in the mercury cell process
a) Electrolysis cell; b) Anolyte tank; c) Vacuum column dechlorinator; d) Cooler; e) Demister; f) Vacuum pump; g) Seal tank;
h) Final dechlorination; i) Saturator; k) Sodium carbonate tank; l) Barium chloride tank; m) Brine reactor; n) Brine filter;
o) Slurry agitation tank; p) Rotary vacuum filter; q) Vacuum pump; r) Brine storage tank; s) Brine supply tank
Figure 22. Cell room: Uhde mercury cells
at 70 – 85 ◦ C. To avoid mercury emissions into
the air, the resaturators are generally closed vessels. The mercury cathode is very sensitive to
poisoning by heavy metals; therefore, a test [90]
has been developed that allows rapid determination of the suitability of any particular salt or
brine.
Chlorine
23
5.3.2. Cell Room
The cells are usually situated in a building
(Fig. 22), although sometimes they are erected
in open air (Fig. 23). Figure 24 shows a bird’s
eye view, and Figure 25 shows a cross section of
cell room. The cells are arranged parallel to each
other so that bus bars and supply lines are kept
short. The cells stand on supporting structures
and are insulated to prevent shorting to the earth.
The transformer and rectifiers are situated at one
end of the room, and the cell service and repair
area is at the opposite end. Ancillary equipment
is installed near the cell room in a spillage containment area.
Figure 24. Mercury cell room (bird’s-eye view, schematic)
a) Cell room; b) Transformer room; c) Rectifier room;
d) Bus bars; e) Turnaround bus bars; f) Service walkways;
g) Ancillary equipment; h) Electrolysis cells; i) Vertical decomposers; k) Cell assembly and maintenance area
Figure 23. Open-air cell room: De Nora cells
Cell floors, gangways, and spillage containment areas are constructed with smooth, sloping
floors so that any mercury can be easily recovered or wash water can be conveniently collected
for treatment. The supply pipes run under the
cells and are connected to them by flexible, insulating connections. The heat given off by the
cells and the decomposers is removed by a ventilation system.
The plant is operated with continuous 24 h/d
supervision and control. An additional day-shift
team carries out anode changes, repairs, and
cleaning.
Figure 25. Mercury cell room (cross section, schematic)
a) Basement floor; b) Floor drains; c) Cell supports with
insulators; d) Supply pipes; e) Cells; f) Decomposers;
g) Service walkways; h) Crane; i) Ridge ventilator;
j) Ventilation air supply; k) Windows/lighting
Occupational Health. Anyone working in
the cell area must undergo regular health checks.
Euro Chlor has prepared a Code of Practice
“Control of Worker Exposure to Mercury in
the Chlor-Alkali Industry” [91]. The U.S. Chlorine Institute has released guidelines “Medical
Surveillance and Hygiene Monitoring Practices
of Worker Exposure to Mercury in the ChlorAlkali Industry” [92]. The U.S. Environmental
Protection Agency has established 18 rules re-
24
Chlorine
lating to cleanliness of the cell room [93]. Adherence to these rules eliminates any danger to
the health of personnel caused by mercury. The
maximum allowable concentration or threshold
limit value (TLV) of mercury in the atmosphere
in Western Europe and in the United States is
between 0.025 and 0.100 mg/m3 .
block, flakes, prills, or powder. For the processes
involved and for uses, see → Sodium Hydroxide.
5.3.3. Treatment of the Products
Chlorine. See Chapter 11.
Hydrogen. The treatment of the hydrogen
gas leaving the decomposer is described in
Chapter 4. It must pass special equipment for
the removal of the traces of mercury before it is
used (see page 25).
Sodium Hydroxide Solution. The great advantage of the mercury cell process is that very
pure sodium hydroxide solution is produced (see
Table 20) at a suitable concentration. The chloride content is only 5 – 50 mg/kg.
Sodium hydroxide from the decomposer usually has a concentration of 50 % and a temperature of 80 – 120 ◦ C. It passes through rubberlined steel pipe work to nickel or Incoloy coolers, where it is cooled to 40 – 60 ◦ C. Any particles of graphite from the decomposer or traces of
mercury are effectively removed by centrifuges,
candle filters, or precoated leaf filters (Fig. 26).
Figure 26. Processing of sodium hydroxide solution from
the amalgam decomposer
a) Vertical decomposer; b) Collection main; c) Collecting
tank; d) Pump; e) Cooler; f) Mercury removal filter
The freezing-point and boiling-point curves
of sodium hydroxide solutions are shown in Figure 27. The phases separating from the solution,
i.e., ice, hydrates, and NaOH, are indicated along
the freezing-point curve. Sodium hydroxide is
supplied to consumers as aqueous solution, solid
Figure 27. Freezing and boiling point curves of sodium
hydroxide solutions
5.3.4. Measurement
The condition of the brine, the cells, and the
products must be continuously and carefully
monitored, since even small deviations from the
correct conditions can increase the hydrogen
concentration in the chlorine. The measuring operations are mostly automatic: critical limits are
chosen and if these limits are exceeded, alarms
are set off.
Cell Operating Conditions. The sodium
concentration in the amalgam is determined at
the cell inlet (max. 0.05 wt %) and outlet (max.
0.45 wt %).
A 20-g sample of the amalgam is reacted with
30 wt % aqueous sulfuric acid in an absorption
pipette. The evolution of 1 cm3 of gas corresponds to 0.01 wt % Na.
A portable analytical and recording apparatus
is available that works electrochemically [94].
If the mercury pump stops, the steel cathode
base of the cell is exposed to electrolyte, and
hydrogen evolves to form an explosive mixture
with the chlorine in the cell.
Chlorine
Failure of the mercury pump or mercury flow
automatically short-circuits the cell. Mercury
flow failure is detected by monitoring the mercury level at the lowest part of the mercury circulation system or in the inlet box, by direct flow
measurement [95], or by loss of pressure at the
pump delivery.
The motors for the mercury pumps, the chlorine absorption plant, and the most vital control equipment are all provided with an emergency power supply, ensuring safe shutdown of
the plant if a power failure occurs. The installation is protected by a complex system of interlocks so that failure of important equipment,
such as the chlorine compressor, shuts down the
rectifiers.
A large number of systems have been developed for the protection of cells from short circuits. Titanium anodes are destroyed by short
circuits and must be raised before any contact
with the amalgam takes place.
The operation of the monitoring system depends on magnetic-field current measurement
for individual anodes [89, 96, 97] or on shunt
measurement of the supply bus bars [98]. Monitoring is achieved by comparison of the anode –
cathode voltage of different cell sections [99] or
by following the conductivity of the brine in the
electrode gap [100].
The signals from the instruments are fed into
central computers or local microprocessors at
each cell [101] that control the anode lifting gear.
The mercury inventory in each cell may be measured by a radioactive tracer technique once a
year without affecting cell operation [102].
Products. The concentration of the sodium
hydroxide solution is determined from its density, and the purity its checked by titration to determine hydroxide, carbonate, and chloride contents. The purity of the water for the decomposer
is determined from its conductivity.
The oxygen content of the hydrogen gas is determined from the magnetic susceptibility, oxygen being paramagnetic.
25
residual emissions to water and air are ecologically acceptable. The mercury in the electrolytic
cells circulates in a closed system. All materials that come into contact with the mercury –
equipment, products, auxiliary chemicals, wash
water, waste gases, other waste materials – may
become contaminated with mercury, and must
be treated before release to the environment or
must be safely deposited. For the exact measurement of these trace amounts and for control of
the effectiveness of the measures to reduce the
emissions, analytical methods have been developed with sensitivities in the microgram region
[103].
Many countries have set legal limits on emissions in waste air and water. The limits on the
products of a chlor-alkali plant may depend on
their end use, e.g., drinking water treatment or
food processing. The sources of contamination
are listed, and means of reducing them are described: Mercury cells are sealed vessels, and the
products are conveyed in closed pipes. The cell
rooms must have smooth joint-free floors with
easily cleaned drainage surfaces and irrigated
collection gutters (see Fig. 25). Spilled mercury
is immediately washed away with water into collecting tanks or sucked up with a vacuum system.
Control of mercury loss is only possible if
the mercury content of all the cells is known exactly. The gravimetric and volumetric methods
formerly used were cumbersome and led to additional mercury emissions, disadvantages that
are avoided by a radioactive tracer method.
Mercury in Products. Hot, moist chlorine
leaving the cell contains small amounts of mercuric chloride. This is almost completely washed
out in the subsequent cooling process and may
be fed back into the brine with the condensate. In
the cooled and dried chlorine gas, there are only
minute traces of mercury: 0.001 – 0.01 mg/kg.
The equilibrium mercury concentration in hydrogen gas is a function of temperature and pressure. The mercury concentration (mg of Hg per
m3 of H2 at 101.325 MPa) increases rapidly with
temperature:
5.3.5. Mercury Emissions (→ Mercury,
Mercury Alloys, and Mercury Compounds,
Chap. 5)
t, ◦ C
0
20
40
60
80
100
c, mg/m3
2.36
14.1
66.1
255
836
2404
Any chlor-alkali plant up to modern technical
standards is not a hazard to the environment. The
Subjecting the hydrogen gas to pressure lowers the mercury content of the resulting prod-
26
Chlorine
uct gas at atmospheric pressure. For example, at
5 ◦C
cHg p = 0.37 Pa kg m−3
cHg = concentration of mercury in hydrogen
gas at atmospheric pressure, mg of Hg
per m3 of H2
p =
pressure to which the hydrogen gas is
subjected, MPa
When the mixture is cooled to 2 – 3 ◦ C, the
mercury concentration is reduced to ca. 3 mg/m3
at standard pressure. This mercury content can
be reduced by compressing and further cooling, adding chlorine to form mercurous chloride
(calomel), which is collected on rock salt or similar material in a packed column, washing with
solutions containing active chlorine, or by adsorption on activated carbon impregnated with
sulfur or sulfuric acid, leaving a mercury concentration in hydrogen of 0.002 – 0.03 mg/m3 .
The highest purity can be achieved by adsorption
on copper/aluminum oxide or silver/zinc oxide,
< 0.001 mg/m3 .
Centrifugation or filtration in candle filters
or in disk filters precoated with charcoal gives
sodium hydroxide solutions containing mercury
concentrations of < 0.05 ppm (mg/kg of 50 %
caustic soda).
The circulating brine contains mercury concentrations of 2 – 20 mg/L. Mercury emissions
from the brine system can occur through losses
of brine into the wastewater, by brine vaporization in the resaturators, or by disposal of
the residues from the brine purification filter.
These emissions are minimal at a chlorine concentration < 30 mg/L, giving a redox potential > 500 mV vs. NHE. Under these conditions
mercury remains dissolved in the brine as a mercury chloride complex even if the brine is alkaline.
Mercury in Wastewater. Mercurycontaining wastewater has several sources:
1) The process, e.g., condensate and wash liquor
from treatment of chlorine, hydrogen, and
brine; stuffing-box rinse water from pumps
and blowers; brine leakages; ion-exchange
eluate from process-water treatment
2) Cell cleaning operations
3) Cleaning of floors, tanks, pipes, and dismantled apparatus
The amount of wastewater can be reduced by
separately disposing of the cooling water and
process water and by feeding the condensate
back into the brine, provided the water balance
allows this. A wastewater rate of 0.3 – 1.0 m3 per
tonne of chlorine is achievable.
There are various methods of making
wastewater suitable for discharge:
1) Chemical removal of mercury by reducing
any compounds to the metal with hydrazine or
sodium borohydride or by precipitating mercuric sulfide with thiourea or sodium sulfide.
The mercury metal or sulfide is then filtered
off.
2) Oxidation of the mercury by chlorine,
hypochlorite, or hydrogen peroxide and adsorption on an ion-exchange medium. Elutriation is done with hydrochloric acid, which is
then used to acidify the brine [104].
The Clean Water Act of 1972 (United States)
demands the use of the “best available technology economically achievable.” Since 1982 each
plant has been limited to a maximum of 0.1 g
of Hg per tonne of chlorine averaged over 30 d
measured at the outlet of the wastewater treatment plant.
In Western Europe, an EC directive has been
issued on the subject of the mercury content of
wastewater from chlor-alkali plants, following
various earlier agreements such as the Rhine protection agreement, the EC guidelines concerning
the protection of natural waters, and the Paris
Convention [105]. This directive requires plants
with circulating brine systems to have a limit of
1.0 g of Hg per tonne of chlorine produced.
Mercury in Process Air. Air from the process, for example, the cell end box ventilation
system, vents from liquid collection tanks (caustic, wastewater), from the vacuum cleaning system, or from the distillation unit for mercurycontaminated wastes can be treated to remove
mercury by the methods used for hydrogen.
Ventilation of the Cell Room. The
heat
produced during electrolysis requires that the air
must be changed 10 – 25 times per hour, depending on the type of building. Mercury spillage can
Chlorine
occur during essential operations involving cells
or decomposers, for example, opening the cells
for anode changing or cleaning, assembling or
dismantling equipment, or replacing defective
pipes. Spillage leads to small losses in the exhaust air owing to the vapor pressure of mercury.
In addition, products that contain mercury, such
as the sodium hydroxide solution, hydrogen, or
process waste air, can escape via faulty seals
in pipes and equipment, leading to emissions.
Closed cell construction and special care in
handling mercury, i.e., good housekeeping, by
adhering to the EPA rules or the Code of Practice “Mercury Housekeeping” [106], keep the
mercury concentrations and, hence, emissions
below the allowable work place concentrations
(MAK and TLV) [107].
Purification of large volumes of waste air
containing mercury in very low concentrations is
not effective. In the United States, the upper limit
for the emission of mercury in process waste air
and hydrogen is 1 kg per day per facility, and
for ventilation air it is 1.3 kg per day per facility [18, p. 372]. In Western Europe the Parcom
Decision 90/3 [108] requires a standard of 2 g
Hg/t of chlorine capacity for emissions to the
atmosphere from existing plants.
Mercury in Residues. Mercury-containing
residues include brine filter slurry, spent decomposer catalyst, discarded cell components,
residues from the purification of products, waste
material from rinsing media, adsorption materials, ion-exchange media, etc. Mercury can be
recovered from these materials by distillation
in closed retorts. The residues after distillation
must be disposed at special sites. Mercury in
safely deposited wastes is not considered an
emission to the environment.
Summary. The continued efforts of all producers have led to a steady decrease in mercury
emissions over the years; for example, in Western Europe from 16 g in 1978 to 2 g Hg per tonne
of chlorine capacity in 1996, as shown in Figure 28 [109]. With this low emission level, the
contribution of the chlor-alkali industry to the total natural and anthropogenic mercury emissions
is less than 0.1 % [111]. Euro Chlor is developing a BAT (best available techniques) for reducing mercury emissions from existing mercurybased plants, the application of which will en-
27
sure that in 2007 no plant emits more than 1.5 g
Hg per tonne of chlorine capacity to air, water,
and products [109].
Figure 28. Mercury emissions from the European chloralkali industry
6. Diaphragm Process
The commercial production of chlorine by electrolytic processes began in Europe and the
United States in the 1890s. Early cells of the
bell-jar type had no diaphragm and relied on the
flow of anolyte toward the cathode to prevent
the hydroxide ion from back-migrating toward
the anode. This method had limited capacity because gas evolution caused mixing and loss of
efficiency. The Griesheim cell, another early design, used porous cement as the diaphragm.
E. A. Le Sueur is credited with the design of
a cell incorporating a percolating asbestos diaphragm, which is the basis for all diaphragm
chlor-alkali cells currently in use. When brine is
caused to flow into the anolyte and subsequently
through the diaphragm into the catholyte, continuous operation with much improved efficiency is obtained. This Le Sueur cell, and the
similar Billiter cell developed in Germany, incorporated a horizontal asbestos sheet as the diaphragm. During the 1920s, the Billiter cell became the most widely used cell in the world; a
few are still in operation.
Following the invention of synthetic graphite,
numerous cells were developed. These fall into
three basic types:
1) Rectangular vertical electrode cells
28
Chlorine
2) Cylindrical vertical electrode cells
3) The vertical electrode bipolar filter press cell
developed by Dow Chemical
In 1913, C. W. Marsh developed a cell with
finger cathodes and side-entering anodes and
cathodes, which greatly increase the electrode
area per unit of floor space. About 1928, Kenneth Stewart of Hooker Chemical (now Occidental Chemical Co.) developed a method of depositing asbestos fibers onto the cathode by immersing the cathode in a slurry of asbestos fibers
and applying a vacuum. All significant installations of diaphragm cells currently in operation
are derived from that development [2].
All diaphragm cells produce cell liquor that
contains ca. 11 wt % caustic soda and 18 wt %
sodium chloride. To market the caustic soda, its
concentration must be increased to 50 %. During the evaporation and cooling processes, the
salt becomes less soluble in the stronger caustic,
and at 50 % NaOH the NaCl concentration is ca.
1 %.
sodium chloride solution (brine) enters the anode compartment and completely covers the anodes and the cathode tubes or fingers. The chlorine leaves the cell through an outlet in the cell
head. The anolyte flows through the diaphragm
into the cathode compartment because of the
difference in liquid level between the two compartments. The catholyte is a solution of sodium
chloride and sodium hydroxide because a portion of the water is converted to hydroxide at
the cathode. The hydrogen produced at the same
time leaves the cell through an outlet on the cathode. The solution of sodium chloride and sodium
hydroxide overflows the cell through a level control pipe on the cathode, and is then commonly
called cell liquor.
6.1. Principles
The principles needed to understand the efficient
operation of the diaphragm process involve the
current efficiency, cell voltage, power consumption, and optimization of the operating conditions [112, 113].
The reaction at the positively charged anode
is the same for all three chlor-alkali processes
2 Cl− → Cl2 +2e−
The reaction at the negatively charged cathode of the diaphragm cell is
2 H2 O+2 e− → H2 +2 OH−
Figure 29 is a cutaway view of a diaphragm
cell that shows the orientation of the various
parts of the cell and the various reactions that
take place. Figure 29 also shows the location of
the diaphragm, which is deposited on the outside of the cathode screen and which separates
the cell into two compartments, one containing
the anodes and one containing the cathode. The
Figure 29. Basic chemical reactions within the cell
a) Anode compartment; b) Cathode compartment;
c) Deposited diaphragm on cathode tubes, rims, and end
screens
Current Efficiency. Current efficiency is defined as the amount of product actually produced
divided by the amount of product that theoretically should have been produced on the basis
of the amount of direct-current electrical energy
input. The current efficiency is never 100 % because of side reactions. The efficiency of a diaphragm cell is usually based on the chlorine
production.
The side reactions that lower the efficiency
are a result of chlorine that enters the catholyte
compartment or of hydroxide ions that enter the
anolyte compartment. The amount of chlorine
that enters the catholyte compartment is small.
Chlorine
The majority of the efficiency losses in a diaphragm cell are due to migration of hydroxide
ions from the catholyte through the diaphragm
and into the anolyte. This back migration takes
place because the negatively charged hydroxide
ions are attracted to the positively charged anodes and because of the hydroxide ion concentration gradient across the diaphragm. This migration of hydroxide ions through the diaphragm
is in equilibrium with the opposing flow of brine
through the diaphragm.
Three factors control the migration of the hydroxide ions into the anolyte:
1) Concentrations of hydroxide and chloride
ions at the cathode side of the diaphragm
2) Flow rate of brine through the diaphragm
3) Condition of the diaphragm
These three factors in dynamic equilibrium
determine the efficiency of the cell.
Factor 1. The higher the concentration of hydroxide ions in the catholyte, the larger the concentration gradient across the diaphragm, and
the higher the probability of hydroxide ions
crossing through the diaphragm. As a result,
changing cell liquor strength strongly affects cell
efficiency. The concentration of chloride ions in
the catholyte also affects cell efficiency because
some of the chloride ions migrate in place of
hydroxide ions.
Factor 2. Decreasing brine flow rate to a cell
increases the conversion of sodium chloride to
sodium hydroxide and raises the hydroxide concentration in the catholyte because of reduced
overflow from the cell. The decreased flow rate
of brine through the diaphragm allows increased
migration of hydroxide ions into the anolyte.
These factors decrease cell efficiency.
Factor 3. The condition of the diaphragm
is extremely important. Nonuniformity in the
diaphragm results in high flow rates of brine
through thin or loosely compacted areas and low
flow rates through thick or compacted areas. In
the areas where there is a low brine flow rate,
back migration of hydroxide is increased.
The degree of inefficiency in a cell is indicated by the two products of the side reactions,
oxygen in the chlorine and sodium chlorate in
the cell liquor. Oxygen in the chlorine gas is the
result of hydroxide ions that migrate through the
diaphragm into the anolyte, where they are oxi-
29
dized:
2 OH− → 1/2 O2 +H2 O+2 e−
Sodium chlorate in the cell liquor is a result
of hydroxide ions that migrate through the diaphragm into the anolyte and react with chlorine
before reaching the anode:
3 Cl2 +6 NaOH → NaClO3 +5 NaCl+3 H2 O
Equations. The simplest equations for calculating cell efficiency are based on the masses
of products produced per unit of electrical input. Theoretically, 1.492 kg of sodium hydroxide and 1.323 kg of chlorine are produced per
kiloampere-hour. It then follows that
Cathode efficiency,% = (kg of NaOH×100) /
(Q×1.492×the number of cells)
Anode efficiency,% = (kg of Cl2 ×100) /
(Q×1.324×the number of cells)
where Q is the quantity of electricity in kA h
Unfortunately, the production of a single cell
cannot be measured with sufficient accuracy to
give meaningful results. To get around this problem, the chlorine industry uses an equation based
on the analysis of the chlorine gas, the cell liquor,
and the anolyte:
%CE = [%Cl2 ×100] /
[%Cl2 +2 (%O2 ) + (%Cl2 ×anox×F ) /cNaOH ]
where
% CE =
anode current efficiency, %
% Cl2 =
percent chlorine in cell gas (air free)
% O2 =
percent oxygen in cell gas (air free)
anox =
oxidizing power of anolyte expressed as grams of NaClO3 per
liter
cNaOH =
NaOH concentration in the cell
liquor, g/L
F =
conversion factor
30
Chlorine
Table 6. Typical voltage distribution a
Current density, kA/m2
Component
voltages, V
1.24
1.55
1.86
2.17
2.48
Anode potential
Cathode potential b
Structure loss c
Brine loss
Diaphragm loss
Intercell bus
1.30
1.12
0.11
0.11
0.24
0.02
1.30
1.13
0.14
0.15
0.31
0.02
1.30
1.15
0.17
0.19
0.36
0.03
1.30
1.16
0.20
0.23
0.41
0.03
1.31
1.17
0.22
0.27
0.47
0.03
Total
2.90
3.05
3.20
3.33
b
3.47
◦
a
OxyTech MDC-55 cell with Modified Diaphragm and expandable anodes. Conditions: anolyte temperature 93 C, anolyte NaCl
concentration 250 g/L, catholyte NaOH concentration 130 g/L.
b
Potential vs. NHE.
c
Includes anode base, anodes, cathode, cathode screens, copper end connectors, and copper side plates.
The denominator is the amount of chlorine
produced plus the amount of chlorine consumed
in the side reactions. This is equivalent to the
amount of chlorine that could have been produced theoretically from the input of current.
The conversion factor F is the product of a volume factor, an electric field factor, and a stoichiometric factor. In practice, it is a function of
cell liquor strength.
The SIX equation is a practical alternative
to the previous equation and is often used with
computers linked to a gas chromatograph and an
automated cell liquor analyzer. The SIX equation is
%CE = [%Cl2 ×100] /
%Cl2 +2 (%O2 ) + %Cl2 ×6×cNaClO3 /cNaOH
This equation also accounts for chlorine lost
to the anolyte. However, it approximates the oxidizing potential of the anolyte with the concentration of chlorate in the cell liquor and assumes
a fixed conversion factor from anolyte concentration to catholyte, namely SIX. The SIX equation approximates the standard equation within
0.5 %.
Cell Voltage. The voltage of a cell is the sum
of five component voltages: anode potential,
cathode potential, cell structure voltage drop, diaphragm voltage drop, and anolyte – catholyte
voltage drop. The anode and cathode potentials
are sums of the reversible voltages, which are
the thermodynamic minimum amounts of work
to cause the reactions to take place, and the overpotentials, which are the additional voltages re-
quired for nonreversible kinetics. The cell structure voltage drop includes the voltage losses in
the cathode, anodes, intercell bus, and all other
connectors in the cell. The sum of the diaphragm
voltage drop and the anolyte – catholyte (brine)
voltage drop is the potential between the electrodes. All of these voltages are functions of current density. Table 6 shows how cell voltage is
strongly affected by cell current. Cell temperature, feed brine NaCl concentration, and the cell
liquor NaOH concentration also affect cell voltage (Table 7), because they affect the conductivity of the solutions between the electrodes. Table
7 clearly shows that current density is the most
important factor.
Table 7. Factors affecting cell voltage: the change in cell voltage
∆U cell divided by the change in four important cell factors
Factor
∆U cell /change in current
density J, mV m2 kA−1
∆U cell /change in cell temperature t cell , mV/◦ C
∆U cell /change in brine
concentration cNaCl , mV L g−1
∆U cell /change in cell liquor
concentration cNaOH , mV L g−1
Modified
cell *
Standard
cell *
450
450
−7.7
−10.1
−0.7
−1.8
0.26
0.6
* OxyTech MDC-55 cell. The modified cell is outfitted with the
Modified Diaphragm and expandable anodes, whereas the standard
cell is outfitted with the standard asbestos diaphragm and the
standard DSA anode. Conditions: anolyte temperature of 93 ◦ C,
anolyte NaCl concentration of 250 g/L, and catholyte NaOH
concentration of 130 g/L.
Excessive brine impurities or other severe operating problems can adversely affect the voltage
of the cell.
Chlorine
Power Consumption. The power consumption of a cell, kWh per tonne of Cl2 , may be
calculated from the cell voltage by the following equation:
Power consumption = Ucell ×756/ε
where
U cell = cell voltage, V
ε =
cell efficiency expressed as a decimal
Optimization. The relationships described
in the preceding paragraphs can be used to determine the optimum economical cell operating
conditions. The optimizations that must be considered are the following:
1) Higher cell liquor caustic strength and lower
steam usage in caustic evaporation versus
lower cell efficiency and higher power consumption
2) Lower current density, lower voltage, and
lower power consumption versus additional
cells and higher capital costs
3) Lower feed brine temperature, thus decreased
steam usage for brine heating, versus higher
cell voltage, lower efficiency, and higher
power consumption
4) High brine pH and reduced acidification costs
versus lower chlorine efficiency, higher power
consumption, and lower product purity
Each diaphragm cell chlorine plant must determine its own optimum conditions for the most
economical operations.
6.2. Diaphragm Cells
Electrolyzers for the production of chlorine and
sodium hydroxide, including both diaphragm
and membrane cells, are classified as either
monopolar or bipolar. The designation does not
refer to the electrochemical reactions that take
place, which of course require two poles or electrodes for all cells, but to the electrolyzer construction or assembly. There are many more
chlor-alkali production facilities with monopolar cells than with bipolar cells.
Bipolar electrolyzers have unit assemblies of
the anode of one cell unit directly connected to
31
the cathode of the next cell unit, thus minimizing
intercell voltage loss. These units are assembled
in series like a filter press, and therefore, the
voltage of an electrolyzer is the sum of the individual cell voltages created by the anode of
one unit, a diaphragm, and the cathode of the
next unit. Bipolar electrolyzers have high voltages and relatively low amperage; therefore, the
cost of electrical rectification is lower per unit of
production capacity. Bipolar electrolyzers either
must be installed in a large number of electrical
circuits or be designed with very large individual cell components. Developers have chosen the
option for large components.
Dow Chemical was the only early developer to have chlorine production needs large
enough to consider the bipolar option [2]. Later,
following the development of the DSA anode,
PPG Industries and Oronzio De Nora Impianti
Elettrochimici designed, and PPG Industries installed Glanor bipolar electrolyzers in a large
complex at Lake Charles, Louisiana [12].
The monopolar electrolyzer is assembled so
that the anodes and cathodes are in parallel.
Therefore, the potential difference of all cells in
the electrolyzer is the same, and the amperage
at any particular current density only depends
on the electrode surface area. A monopolar electrolyzer has low voltage and high amperage. The
highest amperage rating of the most common
modern monopolar cells is ca. 150 kA. Because
a monopolar electrolyzer has a voltage of only
3 – 4 V, circuits of up to 200 electrolyzers have
been constructed, producing 900 t of chlorine
per day.
Diaphragms. The earliest asbestos diaphragms were made of sheets of asbestos paper.
Asbestos was chosen because of its good chemical stability and its ion-exchange properties.
Asbestos has been relatively inexpensive, since
it is a relatively abundant natural material that
was already being mined and processed for other
industrial purposes, such as insulation.
The deposited asbestos diaphragm developed
by Hooker Chemical in 1928 was the most common diaphragm until 1971, when what is now
OxyTech Systems developed the Modified Diaphragm. The Modified Diaphragm is a mixture of asbestos and a fibrous fluorocarbon polymer [114]. The polymer stabilizes the asbestos,
32
Chlorine
which in itself lowers cell voltage and also allows for the use of the expandable DSA anode
[115]. In its various formulations, the Modified
Diaphragm is the most common diaphragm. The
Modified Diaphragm still contains a minimum
of 75 % asbestos.
Environmental concern over the use of asbestos has increased. France, Saudi Arabia, and
Norway have banned the use of asbestos as a separator in chlorine cells. These nations allowed
local chlorine producers several years to install
non-asbestos replacement separators in existing
diaphragm cells or to replace the cells themselves with membrane cells. Other European nations issue permits for the continued use of asbestos until an agreed upon BAT (best available
technology) for diaphragm cells is established.
There is also concern in the chlor-alkaliindustry for the future supply of asbestos, as
most of the North American mines producing
the grades of asbestos previously used for chlorine cell diaphragms have closed. At present the
most common source of asbestos for chlorine
cell diaphragms is Zimbabwe. In addition asbestos disposal costs and regulation have continued to increase. All of theses factors have led the
chlor-alkali industry to consider non-asbestos
diaphragm technology. Three non-asbestos diaphragm systems are commercially available today.
Chloralp Asbestos-Free Technology. Chloralp, a joint venture of Rhône-Poulenc and
La Roche, has developed an asbestos-free diaphragm based on the built-in activation concept. The Chloralp separator is made of two
vacuum-deposited layers:
1) The first layer, known as the precathode, is
a conductive mat of carbon fibers, containing an electrocatalytic powder to decrease the
cathode overpotential
2) The second layer is the diaphragm itself, in
which PTFE and inorganic materials have replaced asbestos
The cathodic activation provides energy savings
from 50 to 150 mV, depending on the plant operating conditions, and the catalyst content in the
mat. Other benefits from the precathode are electrocatalytic destruction of chlorates and lower
hydrogen contents in the chlorine after plant
shutdowns. The diagraphm is deposited on the
precathode by standard vacuum techniques. To
control the diaphragm porosity, a pore-forming
agent is incorporated in the slurry. Optimized
porosity leads to lower cell voltage and higher
current efficiency.
The technology is currently evaluated on
OxyTech HC-3B electrolyzers at the Chloralp
Pont-De-Claix facility. Compared to polymermodified diaphragms, Chloralp asbestos-free
separators provide a 150 mV saving on voltage
and a 2 % improvement in current efficiency.
Combined benefits from the precathode and the
diaphragm, as well as an extended lifetime,
could lead in the near future to full conversion
of the 240 000 t/a Pont-De-Claix plant.
OxyTech Polyramix Diaphragm. OxyTech
Systems has developed and commercialized a
synthetic non-asbestos diaphragm called the
Polyramix diaphragm. The Polyramix fiber is
composed of a PTFE [poly(tetraflouroethylene)]
fibrid base with zirconium oxide ceramic particles embedded in and on the fiber. The Polyramix diaphragm is vacuum deposited onto the
cathode and then baked in an oven to fuse
the fibers together. The process is very similar to OxyTech’s widely used Modified Diaphragms. The Vulcan Chemicals plant in Geismar Louisiana was fully converted to the use
of Polyramix diaphragms in 1993. Most major
diaphragm cell chlorine plants have operated
from 2 to 40 Polyramix diaphragm cells. The
longest life cell with a Polyramix diaphragms
has been in operation for over ten years [116].
PPG Industries Tephram Diaphragm.
PPG Industries has developed and commercialized the Tephram diaphragm as its entry into the
non-asbestos diaphragm market. Major goals of
PPG’s non-asbestos program were to produce
a diaphragm that deposited and operated similarly to asbestos, with longer diaphragm life
and lower power consumption. This technology
utilized vacuum deposition to produce a base
diaphragm composed of PTFE and a topcoat
from a slurry of inorganic particulate materials.
Dopants are priodically added to the anolyte during cell operation to adjust the diaphragm permeability to maintain or improve cell operation.
This diaphragm has been successfully tested of
the major monopolar cell technologies, as well
Chlorine
as on PPG’s Glanor bipolar cells. An active program is in place to operate diaphragms on a trial
basis at various producer sites, allowing for sitespecific evaluation of the overall economics of
the diaphragm.
6.2.1. Dow Cell [2, 117, 118]
The Dow Chemical Company is the largest
chlor-alkali producer, accounting for one-third
of the U.S. production and one-fifth of the world
capacity. Because Dow’s production capacity is
large and concentrated in a few sites, Dow’s cell
development followed a different path than other
chlor-alkali technology developers. Dow uses its
own cell design of the filter press bipolar type.
Dow has operated filter press cells for over 90
years. Dow cell development occurred in several
stages, characterized by simple rugged construction and relatively inexpensive materials.
33
The cell employs vertical DSA coated titanium anodes, vertical cathodes of woven wire
mesh bolted to a perforated steel backplate,
and a vacuum-deposited modified asbestos diaphragm. A single bipolar element may have
100 m2 of both anode and cathode active area.
The anode of one element is connected to the
cathode of the next by copper spring clips. This
connection is immersed in the cell liquor during operation. Figures 30 B and 31 show these
internal cell parts.
Figure 31. Dow diaphragm cell, section view
a) Perforated steel backplate; b) Cathode pocket;
c) Asbestos diaphragm; d) DSA anode; e) Copper
backplate; f) Titanium backplate
Figure 30. Dow diaphragm cell
A) Six-cell series
B) Internal cell parts: a) Cathode element; b) Cathode
pocket elements; c) Copper spring clips; d) Perforated steel
backplate; e) Brine inlet; f) Chlorine outlets; g) Copper
backplate; h) Titanium backplate; i) Anode element
Dow operates at lower current densities than
others in the chlor-alkali industry. The electrolyzers are normally operated with 50 or more
cells in one unit or series. One electrical circuit may consist of only two of these electrolyzers. Figure 30 A shows a view of six electrolytic
cells.
Treated saturated brine is fed to the anolyte
compartment, where it percolates through the
diaphragm into the catholyte chamber. The percolation rate is controlled by maintaining a level
of anolyte to establish a positive, adjustable hy-
34
Chlorine
drostatic head. The optimum rate of brine flow
usually results in the decomposition of ca. 50
% of the incoming NaCl, so that the cell liquor
is a solution containing 8 – 12 wt % NaOH and
12 – 18 wt % NaCl.
The Dow diaphragm cell, optimized for low
current density, consumes less electrical energy
per unit of production than the rest of the industry. The cell voltage at these low current densities
is only 300 – 400 mV above the decomposition
potential of the cell. However, Dow has a larger
investment in the electrolyzers, especially anodes.
The electrolyzers are operated at ca. 80 ◦ C,
lower than the 95 ◦ C typical of other types of
cells. This lower operating temperature allows
cell construction with less expensive materials, such as vinyl ester resins and other plastics
[117]. Operating data have not been published.
6.2.2. Glanor Electrolyzer [12, 119 – 121]
Glanor bipolar electrolyzers are a joint development of PPG Industries and Oronzio De
Nora Impianti Elettrochimici S.p.A. The Glanor
electrolyzer consists of several bipolar cells
clamped between two end electrode assemblies
by means of tie rods, thereby forming a filter press type electrolyzer (Fig. 32). The electrolyzer is equipped with DSA titanium anodes.
Each electrolyzer normally consists of 11 or 12
cells. A lower number of cells can, however, be
assembled in one electrolyzer. The Glanor electrolyzer was especially designed for large chloralkali plants.
The current is fed into the electrolyzer by
means of anodic and cathodic end elements. The
anodic compartment of each cell is connected
to an independent brine feed tank by means of
flanged connections.
Chlorine gas leaves each cell from the top
through the brine feed tank and then passes to
the cell room collection system. Hydrogen gas
leaves from the top of the cathodic compartment
of each cell, while the catholyte liquor leaves
from the bottom through an adjustable level connection.
Figure 32. Glanor bipolar electrolyzer
a) Disengaging tank; b) Chlorine outlet; c) Hydrogen outlet; d) Bipolar element; e) Brine inlet; f) Cell liquor trough;
g) Cell liquor outlet
Figure 33. Glanor bipolar electrolyzer type V-1144
The V-1144 electrolyzer (Fig. 33) was the
first commercial unit, and eight plants utilize this model. The second generation is the
V-1161 electrolyzer, which employs Modified
Diaphragms, narrower electrode gaps, lower
current density, and DSA anodes to achieve
lower power consumption than the V-1144 electrolyzer.
The operating characteristics of the Glanor
electrolyzers are shown in Table 8.
Chlorine
Table 8. Glanor bipolar diaphragm electrolyzers: design and
operating characteristics
Item
Model
V-1144
Model
V-1161
Cells per electrolyzer
Active anode area per cell, m2
Electrode gap, mm
Current load, kA
Current density at 72 kA, kA/m2
Cell voltage, V
Current efficiency, %
Power consumption (d.c.),
kWh/t *
Anode gas composition
(alkaline brine)
Cl2 , %
O2 , %
H2 , %
CO2 , %
Cell liquor
NaOH, g/L
NaClO3 , %
Production per electrolyzer
Chlorine, t/d **
NaOH, t/d
11
35
11
72
2.05
3.50
95 – 96
2500
11
49
6
72
1.47
3.08
95 – 96
2200
97.3 – 98.0
1.5 – 2.2
<0.1
0.4
97.0 – 98.0
1.5 – 2.2
<0.1
0.4
35
with both ends open, extending across the cell,
as the circulation space requirement was satisfied by the change from solid graphite anodes to
the open DSA anodes (Fig. 34).
Figure 34. OxyTEch type H-4 cell
135 – 145 135 – 145
0.03 – 0.15 0.03 – 0.15
26.7
29.8
26.7
29.8
* Per short ton of chlorine.
** Short tons.
6.2.3. OxyTech “Hooker” Cells [2, 12, 122,
123]
The first commercialized deposited asbestos diaphragm cell was the Hooker type S-1 monopolar cell, introduced in 1929. The basic design featured vertical graphite anode plates connected to
a copper bus bar and a cathode with woven steel
wire cloth or perforated steel fingers between the
anodes. The cathode held vacuum-deposited asbestos fiber diaphragms that separated the anode
and cathode compartments. The cathode fingers
did not extend completely across the cell, but
left a central circulation space. In the following
40 years, a family of S series cells with similar
characteristics evolved, with over 12 000 having
been installed in licensed plants.
In 1973, a new H series of monopolar cells
was introduced. They incorporated the use of
DSA anodes, which had been developed and
commercialized in the late 1960s. These cells
have significant voltage savings over the S series,
thus allowing increases in cell capacity without
corresponding increases in rectification capacity. The H series also incorporate cathode tubes
Figure 35. OxyTech/Uhde HU-type cells
a) Cell bottom; b) Cathode; c) Anode; d) Cell cover; e) Bus
bars; f) Brine level gauge; g) Brine flow meter; h) Bypass
switch
Table 9 is a summary of operating characteristics and current densities of the H-series cells
currently available for license.
6.2.4. HU Monopolar Cells [123]
The HU type cells were a joint development
of Hooker (now OxyTech Systems) and Uhde.
The HU-type electrolyzer (Fig. 35) is rectangular, not cubic, and is narrow in the direction of
current flow, since anodes are arranged in a sin-
36
Chlorine
Table 9. OxyTech Systems Hooker H-series diaphragm cells: design and operating characteristics
H-2A
Operating current, A
Anode area, m2
in.2
Current density, A/m2
A/in.2
Cell voltage, V
Approximate cell
dimensions, m
Diaphragm life, days
Anode life, years
Operating NaOH
concentration, g/L
%
H-4
80 000
36.16
56 050
2212
1.43
3.44
1.87×2.66
Current Efficiency, %
Chlorine output,
metric ton/day
short ton/day
Caustic soda output,
metric ton/day
short ton/day
150 000
64.52
100 000
2325
1.50
3.44
2.58×3.11
300 – 500
5–7
5–7
300 – 500
140
11.35
160
12.89
140
11.33
160
12.87
96.4
94.6
96.6
94.9
2.45
2.70
2.41
2.65
4.60
5.07
4.52
4.98
2.76
3.05
2.71
2.99
5.19
5.72
5.10
5.62
Table 10. HU series diaphragm cells: design and operating characteristics
Item
Number of anodes
Anode surface area, m2
Load, kA
Cl2 production, t/d
NaOH (100 %) production, t/d
H2 production, kg/d
Cell length, m
Distance, cell-to-cell, m
Cell type
HU 24
HU 30
HU 36
HU 42
HU 48
HU 54
HU 60
24
20.6
30 – 45
0.90 – 1.36
1.01 – 1.54
25 – 39
2.1
1.5
30
25.8
40 – 60
1.19 – 1.82
1.35 – 2.05
34 – 52
2.6
1.5
36
31.0
50 – 70
1.49 – 2.12
1.68 – 2.39
42 – 60
3.0
1.5
42
36.1
55 – 85
1.64 – 2.58
1.85 – 2.91
47 – 73
3.5
1.5
48
41.3
60 – 95
1.79 – 2.88
2.02 – 3.25
51 – 82
3.9
1.5
54
46.4
70 – 105
2.09 – 3.18
2.36 – 3.59
59 – 91
4.4
1.5
60
51.6
80 – 120
2.39 – 3.64
2.69 – 4.10
68 – 103
4.8
1.5
gle row. The cathode is long and narrow; consequently, the current density is lower through the
cathode shell. The long, narrow cathode fabrication lends itself to closer anode – cathode tolerances and spacing. Copper on and around the
cathode shell has been eliminated. Another advantage of the long, narrow design is shorter
electrolysis current paths through the cell room,
resulting in savings in piping and other materials. The HU-type cell incorporates a Modified
Diaphragm.
A further novelty of the HU cell system is the
design and arrangement of the bypass switch.
The HU switch is installed underneath, not next
to, the circuit of cells. This is accomplished by
raising the cells from the floor, similar to mer-
cury cells, creating a second operating floor. The
interconnecting bus bars are flexible and are distributed over the entire length of the cell. The
HU cell design incorporates a bus bar for each
individual anode. This, as well as the elevation
of the cell from the floor below, which allows
access, enables connection of facilities for monitoring the current flowing through each anode.
During operation of the bypass switch, connection is made for each individual anode, and no
additional contact bus bars are required.
The HU-type cells are offered to cover 30 –
150 kA. All of the different cell types are
equipped with cathodes and anodes of identical height and width. The only basic difference
between the various cell models is the number of
Chlorine
elements and consequently the length of the cell
(Table 10). Cell voltage and power consumption
per tonne of chlorine, identical for all cell types,
are shown in Table 11 for the specific current
loads of 1.5 and 2.3 kA/m2 .
Table 11. HU series diaphragm cells: specific load, cell voltage,
and power consumption
Specific load, kA/m2
Cell voltage, V
Power consumption
(d.c., average), kWh/t*
1.5
2.3
3.12
3.41
2500
2700
* Per tonne of chlorine.
37
Copper connectors attached at the ends of the
bonded copper side plates complete the encompassing of the cathode with copper. Anodes are
connected to a copper patented cell base, which
is protected from the anolyte by a rubber cover or
a titanium base cover (TIBAC) [125]. Orientation of the cathode tubes is parallel to the cell circuit, the opposite of a Hooker-type cell. This arrangement accommodates thermal expansion of
the cell and circuit without changing the anodeto-cathode alignment.
The combination of the Modified Diaphragm
and expandable DSA anodes reduces power consumption by 10 – 15 % from that of regular asbestos diaphragms and standard, fixed DSA anodes [126]. Table 12 presents performance data
for the two most common MDC cell sizes [124].
The OxyTech MDC-29 is shown in Fig. 38. The
licensed chlorine capacity of OxyTech cells now
exceeds 20 000 t/d.
6.3. Operation
Figure 36. OxyTech Systems MDC cells
a) Brine feed rotometer; b) Head sight glass; c) Cell
head; d) Cathode assembly; e) Tube sheet; f) Grid plate;
g) Cathode tube; h) Grid protector; i) DSA expandable anode
6.2.5. OxyTech MDC Cells [12, 124]
OxyTech Systems manufactures and licenses
the MDC series of monopolar diaphragm cells
(Fig. 36). The MDC cells feature woven steel
wire cathode screen tubes open at both ends,
which are welded into thick steel tube sheets at
each end. The tubes, tube sheets, and the outer
steel cathode shell form the catholyte chamber
of the cell (Fig. 37). Copper is bonded, rather
than welded, to the rectangular cathode shell on
the two long sides parallel to the tube sheets.
The process description in this section is intended to provide an overview of typical diaphragm cell process areas. A general block diagram for a diaphragm cell facility is shown in
Figure 9. Included on the drawing are many process areas that may be optional, depending on
the design of the plant and its end products. The
operation of a cell room may be broken down
into six areas: the two incoming systems, brine
and electrical; the cells; and the three outgoing systems, chlorine, hydrogen, and cell liquor.
Some of these are essentially the same for all
three chlor-alkali processes and are described
in Chapter 4 — the brine system (general), the
electrical system, and the hydrogen system. The
treatment of the chlorine is the subject of Chapter 11. Only aspects that are reasonably specific
to the diaphragm cell process are described in
this section.
6.3.1. Brine System
Most commonly, diaphragm cells are supplied
with well brine on a once-through basis. The
treated well brine flows to the treated brine storage tanks, which usually have 12-h capacity.
From there the brine is fed to the cell room. The
38
Chlorine
Table 12. OxyTech Systems MDC cells: operating capacities and characteristics
Item
Model number and operating range, kA
MDC-29
MDC-55
35 to
80
75 to
150
Chlorine capacity,
metric ton/day
short ton/day
1.05
1.16
2.41
2.66
2.33
2.48
4.53
5.00
Caustic capacity,
metric ton/day
short ton/day
1.21
1.33
2.76
3.04
2.59
2.85
5.18
5.70
Hydrogen capacity,
m3 /day
cubic feet/day
335
11 830
765
27 010
720
25 420
1435
50 670
Current density,
kA/m2
A/in.2
1.21
0.78
2.76
1.78
1.37
0.88
2.74
1.76
Cell voltage, V a
steel cathode
activated cathode
2.90
2.80
3.62
3.51
3.00
2.90
3.62
3.51
Power consumed
(d.c., steel
cathode) b ,
kW h/t
kW h/short ton
2310
2100
2876
2610
2390
2175
2870
2610
Power consumed
(d.c., activated
cathode) b ,
kW h/t
kW h/short ton
2230
2025
2786
2530
2310
2100
2780
2530
Diaphragm life, years
Anode life, years
Cathode life, years
1–2
8 – 10
10 – 15
0.5 – 1.0
5–8
10 – 15
1–2
8 – 10
10 – 15
0.5 – 1.0
5–8
10 – 15
Distance between
cells c ,
m
inches
a
b
c
1.60
63
2.13
84
Cell voltage includes loss in intercell bus.
Power consumed per ton (metric or short) of chlorine produced.
Distance centerline-to-centerline and side-by-side with bus connecting.
flow to each individual electrolyzer is controlled
by a rotameter. If the flow of brine to the cells is
suddenly disrupted by failure of the brine feed
pump, the rectifiers automatically shut down
since an inadequate supply of brine to the cells
is potentially dangerous. The specifications for
brine for diaphragm cells are given in Table 13.
A brine recovery lagoon is usually available
to handle any major upsets in the brine system.
Brine sludges or out-of-spec brine can be sent
to the lagoon. Supernatant clear brine can be recovered from the lagoon.
In most cases, operation with acidic brine is
preferred because of the reduced amount of sidereaction products in the chlorine and the cell
liquor.
Chlorine
39
Figure 37. Exploded view of an OxyTech MDC-55 cathode
a) End plate; b) Rim screen; c) Side screens; d) Tube sheet; e) Full cathode tube; f) Half-cathode or end tube; g) Side
plate; h) Lifting lug; i) Punched and coined stiffener strap; j) Bosses; k) End plate, operating aisle end; l) Hydrogen outlet;
m) Connector bar; n) Caustic outlet; o) Clip angles; p) Grid bar, connector side; q) Side plate
6.3.2. Cell Room
Typical cell rooms are shown in Figures 39
(bipolar cells) and 40 (monopolar cells).
Figure 39. Cell room: bipolar PPG Industries/De Nora
Glanor cells
Figure 38. OxyTech MDC-29 with the author (1971)
A cell in normal operation requires little attention. The critical requirement is that the brine
flow rate is sufficient to maintain an anolyte level
above the cathode.
40
Chlorine
Table 13. Typical brine feed specifications for diaphragm cells
Parameter
Specification
NaCl
pH
Hardness (Ca2+ + Mg2+ )
Magnesium
Sodium sulfate (Na2 SO4 )
Organics
Manganese
Barium
Nickel
Iron
Silicon
Cobalt
Mercury
Phosphate
320 g/L
2.5 – 3.5
<5 ppm
<0.4 ppm
<5 g/L
<1 ppm
<0.01 ppm
<0.01 ppm
<0.1 ppm
<0.5 ppm
<15 ppm
<0.02 ppm
<1 ppm
<1 ppm
Figure 40. Cell room: monopolar OxyTech H-4 cells
Under no circumstances should a cell be operated with an inadequate or excessive anolyte
level. Operation with the anolyte level not visible
in a sight glass is unsafe. At least one operator
should be in the cell room at all times. The cell
room operator should inspect the anolyte level
and brine flow to each cell at least once per hour.
Any change in the anolyte level or brine flow rate
should be investigated.
As the cell ages, the diaphragm will undergo
changes in porosity because of the following:
1)
2)
3)
4)
Electrolysis effect
Brine impurities
Upsets in operation
Gradual wear of the diaphragm
A change in porosity may necessitate a
change in brine flow rate. If the increase in porosity is severe, the cell may be replaced or doped
with an asbestos slurry or inorganic salts.
Impurities in the brine often lead to decreased
porosity. Decreased porosity can be offset to
some extent by increasing the anolyte level and,
if necessary, by lowering the catholyte level.
A cell operated with the anolyte level at the
maximum value and the lowest catholyte level
is called a sleeper. To gain additional diaphragm
life after a cell has entered the sleeper position,
the brine flow rate must be decreased below normal. This is not normally a recommended practice because current efficiencies of these cells
are usually low.
For safe operation of diaphragm cells, the
header pressures must be maintained at the
proper values.
The chlorine header should be maintained at
positive pressure to permit detection and correction of any chlorine piping leaks. The hydrogen
header is also maintained at a positive pressure
to avoid pulling air into the hydrogen, creating a
potentially explosive mixture. The brine header
pressure should be maintained to give the desired
caustic concentration in the cell liquor. Normal
practice is to adjust individual brine feed valves
so that each cell receives the correct brine flow
rate.
Load changes must be smooth to avoid fluctuations in the header pressures and detrimental
effects on the diaphragms.
The brine feed rate to each cell should be increased to the new rate before circuit amperage is
increased. The brine feed rate to each cell should
be decreased immediately after amperage is decreased. During any period of operation when
brine flow rates are being changed, extra attention should be given to the anolyte levels of the
cells. Adjustment of the brine feed temperature
may also be necessary when a load change occurs.
Chlorine
41
Figure 41. Process flow diagram: triple-effect caustic evaporator
a) First-effect vapor body; b) First-effect heat exchanger; c) Second-effect vapor body; d) Second-effect heat exchanger;
e) Second-effect forwarding pumps; f) 50 % caustic transfer pumps; g) Third-effect vapor body; h) Third-effect heat exchanger;
i) Third-effect forwarding pump; j) Barometric condenser; k) First-stage ejectors; l) Intercondenser; m) Second-stage ejector;
n) Liquor flash tank
6.3.3. Diaphragm Aging
Of all the cell components, the diaphragm usually has the shortest life. The ability of a diaphragm to resist the back migration of hydroxide slowly becomes impaired with service life.
The performance of the diaphragm deteriorates
for the following reasons:
1) Chemical attack
2) Brine impurities
3) Unsteady operating conditions
The major reason for the deterioration is
chemical attack on the asbestos by the alkaline
catholyte and acidic anolyte. The rate of chemical attack can be minimized and diaphragm life
maximized by careful operation of the cell. The
most important situations to avoid are high concentrations of brine impurities and unsteady operating conditions. High brine impurities cause
plugging of the diaphragm with insoluble hydroxides, which reduce the diaphragm’s separation ability. The most common harmful impurities are calcium, magnesium, iron, nickel,
silicates, aluminum, manganese, and barium.
Unsteady operation, such as electrical load
changes, cell liquor strength changes, changes
in brine concentration or pH, gas-pressure fluctuations, and shutdowns, change the pH of the
various regions of the diaphragm, thus accelerat-
ing chemical attack on the asbestos. Diaphragm
cell plant operators should strive to minimize
these changes.
The real importance of the equations in Section 6.1 is as an aid in deciding when the diaphragms should be replaced.
6.3.4. Treatment of the Products
Chlorine. See Chapter 11.
Hydrogen. See Chapter 4.
Sodium Hydroxide Solution. The hydroxide produced at the cathode is associated with
sodium ions and water to form a 10 – 12 wt %
sodium hydroxide solution leaving the electrolytic cell. This cell liquor also contains
18 wt % unreacted sodium chloride.
Most large modern chlor-alkali plants have or
will soon have an associated cogeneration power
plant. In these facilities, the caustic evaporators
are an important use for the byproduct steam.
Modern diaphragm cell plants use tripleeffect evaporators and, in many cases,
quadruple-effect evaporators.
42
Chlorine
Figure 42. Caustic purification system
a) 50 % caustic feed tank; b) 50 % caustic feed pumps; c) Caustic feed preheater; d) Ammonia feed pumps; e) Ammonia
feed preheater; f) Extractor; g) Trim heater; h) Ammonia subcooler; i) Stripper condenser; j) Anhydrous ammonia storage
tank; k) Primary flash tank; l) Evaporator reboiler; m) Evaporator; n) Caustic product transfer pumps; o) Purified caustic
product cooler; p) Purified caustic storage tank; q) Ammonia stripper; r) Purified caustic transfer pumps; t) Overheads condenser; u) Evaporator; v) Evaporator vacuum pump; w) Aqueous storage ammonia tank; x) Ammonia scrubber; y) Scrubber
condenser; z) Ammonia recirculating pump; aa) Ammonia recycle pump
Caustic Soda Evaporation. A flow diagram
for a typical triple-effect caustic soda evaporator is shown in Figure 41. The evaporator is of
the backward-feed design and concentrates 10 –
11.3 wt % NaOH cell liquor to 50 wt % NaOH.
Liquor flows from the third to the second to the
first effect and from the first effect to the liquor
flash tank. A cyclone is used for each effect to
utilize the pressure drop across the circulating
pump to clarify the transfer liquor.
The salt precipitated in the liquor flash tank
is isolated from the rest of the salt precipitated
in the evaporator and used as seed crystals in
the cooling system to help diminish coil scaling
and supersaturation of the product liquor with
sodium chloride. The sodium chloride and triple
salt (NaOH – NaCl – Na2 SO4 ) precipitated in
the liquor flash tank and cooling system is removed from the cooled product liquor with centrifuges. The salt precipitated in the three effects
flows countercurrent to the liquor flow so that all
of the salt is discharged from the last effect, the
effect that has the coldest liquor and the lowest
caustic soda concentration.
Two-stage steam-jet air ejectors with a common intercondenser are used to maintain vacuum in the evaporator. In the caustic cooling
system, agitated tanks are used to cool the slurry
discharged from the liquor flash tank. The slurry
flows from the cooling system to the centrifuge
feed tank from where it is pumped into centrifuges. Salt discharged from the centrifuges
drops into the evaporator feed tank, where it
is dissolved in cell liquor. The 50 wt % NaOH
concentrate liquor, which flows by gravity to
the pressure filter feed tank, contains ca. 1.0 –
1.5 wt % dissolved NaCl and ca. 0.1 wt % crystalline NaCl.
The liquor is pumped from the pressure filter
feed tank into the pressure leaf filters, where the
remaining traces of salt crystals are removed.
The product caustic flows by gravity to the filtered product tank and then is pumped to storage.
Salt removed in the pressure filters is reslurried
with cell liquor and pumped to the evaporator
feed tank via the filter backwash pump.
The salt discharged from the centrifuges
drops into the leaching tank, where it is reslurried with condensate and recycled brine from
the Glauber’s salt (Na2 SO4 · 10 H2 O) crystallizer. Concentrate from the pusher centrifuges
flows by gravity into the evaporator feed tank.
The product salt is discharged from a cyclone
into the salt reslurry tank. The overflow from
the cyclone is returned to the leaching tank. The
product salt is diluted with brine and pumped to
the resaturator tank.
Brine containing the dissolved sodium sulfate is separated from the salt crystals in a cyclone. The underflow returns to the leaching
tank. The overflow is collected in the feed tank
for the Glauber’s salt crystallizer. Sodium sulfate
is crystallized from the liquor in a continuous
vacuum cooled crystallizer.
Mother liquor removed from the crystallizer
is pumped under level control to the brine tank.
Slurry discharged from the crystallizer is thickened to ca. 50 wt %. The liquor from the thick-
Chlorine
ener is collected in the brine tank and pumped
back to the leaching tank. The thickened slurry is
redissolved in the Glauber’s salt dissolving tank
and pumped to a waste treatment system.
Caustic Purification [127]. Diaphragm-cell
chlor-alkali producers requiring higher purity
caustic than that produced by the diaphragm process can use caustic purification or DH process
(Fig. 42). Salt removal in the purification unit is
effected by contacting the 50 wt % caustic with
anhydrous liquid ammonia under pressure sufficiently high to maintain all materials in the liquid
state.
The liquid ammonia absorbs salt, chlorate,
carbonate, water, and some caustic. It is then
stripped, concentrated, and returned to the extraction process. The concentrated caustic leaving the extractor is stripped free of ammonia,
which is recovered, concentrated, and recirculated. Typical purities before and after caustic
purification are shown in Table 20. This process is offered for license by PPG Industries and
OxyTech Systems.
In addition, producers and users of
diaphragm-cell caustic may wish to reduce
metal impurities by utilizing the porous cathode cell process (PPG Industries) [128]. The
process consists of an electrolysis cell with
porous nonmetallic cathodes. The caustic soda
(50 wt %) is freed from iron, nickel, lead, and
copper, which are deposited on the cathode. The
cell must be regenerated periodically with water
and hydrochloric acid. Typical feed and product
analyses based on anhydrous NaOH are
Metal
Iron
Nickel
Lead
Copper
43
ampere load on each circuit
voltage for each circuit
chlorine header pressure
hydrogen header pressure
brine header pressure or flow rate
brine temperature
brine pH
cell liquor temperature
Samples of brine should be taken every 4 h
and combined into a daily composite. In addition, samples of cell liquor should be taken from
each cell string, the sodium hydroxide content
analyzed, and the temperature taken every 4 h.
A daily composite should be made and samples
should be analyzed by the laboratory for the following:
NaOH content
NaCl content
salt : caustic ratio
NaClO3 content
NaOCl content
Fe content
average temperature
specific gravity at 25 ◦ C
Chlorine gas from each cell circuit should be
analyzed for chlorine and hydrogen content at
least twice each 8-h shift. Each day a complete
analysis of the chlorine header gas should be
made. Additions or extensions to list may be dictated by plant operation.
Each plant must develop a procedure for taking individual cell data so that individual cells
may be scheduled for renewal. The following is
a minimum schedule:
Content, ppm
Feed
Product
10.0
3.0
4.0
0.2
2.0
0.2
0.4
0.1
weekly
voltage and cell liquor composition
(NaOH, NaCl, NaClO3 )
monthly
chlorine composition (Cl2 , O2 , H2 ,
CO2 , N2 )
6.3.5. Measurement
7. Membrane Process
Recorded data is an important tool for determining the operating condition of the plant and diagnosing problems.
The following should be recorded continuously or hourly:
In the membrane process, the anolyte and
catholyte are separated by a cation-exchange
membrane that selectively transmits sodium ions
but supresses the migration of hydroxyl ions
from the catholyte into the anolyte. A strong
caustic soda solution with a very low sodium
44
Chlorine
chloride content can be obtained as the catholyte
efflux. The advantages of the membrane process
are its energy efficiency and its ability to produce
caustic soda of high quality, with almost no impact on the environment. Depending on the particular design, membrane sizes from 0.2 to 5 m2 .
The production capacity of an electrolyzer can
be up to 90 t/d NaOH (100 %).
The process was started in the early 1970s
with development of the perfluorosulfonate
membrane Nafion by DuPont [129]. In 1975, a
perfluorocarboxylate membrane capable of producing 35 wt % caustic soda became available,
from Asahi Glass in Japan [130]. In 1978 the
first two-layer membrane was developed, with
low electrical resistance and high current efficiency [131].
The industrial success of the membrane process started in Japan, where the abolition of the
mercury process on environmental grounds had
been promoted by the government. Today the
membrane process is the state of the art process
for producing chlorine and caustic soda or potassium hydroxide.
The production capacity of chlor-alkali plants
using the membrane process reached about 21 %
of total world production capacity in 1995 and
is predicted to increase to about 28 % in 2001
[132].
The anolyte is discharged from the cell.
The electric field causes hydrated sodium ions
to migrate through the membrane into the
catholyte. In the cathode compartment, hydrogen is evolved at the cathode, leaving hydroxyl
ions, which together with permeating sodium
ions constitute the caustic soda:
2 H2 O+2 e− → H2 +2 OH−
Na+ +OH− NaOH
Liquid and gaseous phases anolyte/Cl2 and
catholyte/H2 can be separated either in the cell
compartment or downstream of the cell outlet.
The chlorine-saturated anolyte is then treated in
a dechlorination unit to recover the dissolved
chlorine.
Membrane.
Structure. The membrane is exposed to chlorine and anolyte on one side and strong caustic
solution on the other side at high temperature
(90 ◦ C). Only ion-exchange membranes made
of perfluoropolymer can withstand such severe
conditions. The ion-exchange groups of the original polymers are in the fluorosulfonate form,
–SO3 F, or the carboxylate form, –COOR.
7.1. Principles
In a membrane cell a cation-exchange membrane separates the anolyte and catholyte, as
shown in Figure 43. Saturated brine is fed into
the anode compartment, where chlorine gas is
evolved at the anode:
Fluorosulfonate form
2 Cl− → Cl2 +2 e−
Carboxylate form m = 0 – 1 n = 1 – 5 R = alkyl
The first membranes that showed significant potential for use in the chlor-alkali process were made of a perfluorosulfonate layer.
These proved to be durable in chlor-alkali cells
but were relatively inefficient. The perfluorocarboxylate polymers, with a lower water content,
Figure 43. Principle of the membrane cell
Chlorine
showed higher selectivity but led to higher electrical resistance and high electrical power consumption [133]. Combining the advantages of
high current efficiency and low electrical resistance a composite membrane (Fig. 44) was developed with a layer containing SO−
3 groups on
the anode side and a layer containing COO−
groups on the cathode side [131, 134]. The perfluorosulfonate layer is thicker than the perfluorocarboxylate one and is the major constituent
of the membrane [133].
45
gen gas can be depressed by selecting an anode
coating with suitable characteristics (see Section
8.1) or by decreasing the pH in the anode compartment by acidifying the inlet brine (Fig. 45).
Figure 45. Oxygen content in chlorine [135]
Figure 44. Membrane structure
Flux through the Membrane. The total flux
through the membrane can be divided in to three
parts [133]:
1) Migration due to electric field
2) Convection
3) Diffusion due to chemical gradients
Migration is the flux of ions through the membrane, driven by electric field. This includes the
desired transfer of sodium ions to the cathode
compartment and the undesired transfer of hydroxyl ions to the anode compartment. The capacity for selective separation of the cation exchange membrane is determined by its repulsive
force for hydroxyl ions. This effect determines
the current efficiency.
The backmigration of hydroxyl ions increases the formation of oxygen, hypochlorite,
and chlorate in the anode compartment and
causes a loss of current efficiency of 3 – 7 % in
caustic soda production. The evolution of oxy-
Convection and diffusion determine the flow
of uncharged compounds and ions through the
membrane. Chloride anions in the catholyte are
excluded by the cation-exchange membrane and
repelled by the electric field, so that the transfer rate of chloride anions from the anolyte is
extremly low. As a result, a caustic soda solution of about 32 – 35 wt % with a salt content
of less than 20 ppm can be obtained. The water transport through the membrane is about 3.5
to 4.5 moles of water per mole of sodium ions,
and can be regarded as the hydration sphere of
the migrating sodium ions. Water flux increases
with decreasing anolyte concentration [136].
Migration, convection and diffusion influence each other, and the resulting flux depends
on membrane type, current density, temperature,
and composition of anolyte and catholyte.
Cell Voltage. The cell voltage of a membrane cell is composed of the following terms:
1) Decomposition voltage
2) Membrane potential between anolyte and
catholyte
3) Electrode overpotentials for chlorine and hydrogen
4) Ohmic drop in the membrane
5) Ohmic drop in the electrolytes
6) Ohmic drop in electrodes and conductors
Term 1. The decomposition voltage of the
chlor-alkali process is about 2.20 V, depending
on temperature, concentration, and pressure.
46
Chlorine
Term 2. This term describes the overpotentials at the surfaces of the membrane. Under standardized operating conditions (3 kA/m2 ,
90 ◦ C, 32 wt % caustic solution), the membrane
potential is approximately 0.08 V.
Term 3. Titanium anodes coated with oxides
of Ir, Ru or Pt are generally used in membrane
cells and lead to a chlorine overvoltage of approximately 0.05 V at 3 kA/m2 . Hydrogen overpotentials of about 0.1 V at 3.0 kA/m2 are attained with activated cathodes. Mainly nickel
substrates are coated by painting and thermal
treatment or galvanic deposition [137, 138].
Coating materials include Ni, Co, Ru and others
(see Section 8.2).
Term 4. The ohmic drop of advanced commercial membranes under standardized operating conditions is about 0.25 – 0.30 V at 3 kA/m2 .
Term 5. To minimize the ohmic drop of an
electrolyte, the gaps between the membrane and
the electrodes are minimized in membrane cells.
However, if the gap is very small, a rise in voltage
is observed due to the entrapment of gas bubbles
between the electrodes and the hydrophobic fluoropolymer membrane.
Term 6. The voltage losses in the electrolysis cell that occur due to unfavorable current
paths along the metallic structure are reduced
by an appropriate design. In modern chlor-alkali
electrolysis cells the typical ohmic drop is 20 –
40 mV at 3 kA/m2 .
Current Efficiency. The current efficiency
(CE) for caustic soda can be obtained either by
directly measuring the quantity of caustic soda
produced or by an anodic balance, i.e., the compositions of the anode gas and the anolyte with
the following equation:
approximately 2 – 4. This leads to a pH gradient
across the membrane cross section. The solubility of impurities, which are always present in
the pure brine, depends on the pH. Therefore,
depending on the type of impurities and on the
pH, precipitation inside the membrane can take
place. This leads to mechanical destruction of
the membrane, which has a irreversible effect
on current efficiency. In addition, the cell voltage rises due to the crystals formed inside the
membrane.
7.2. Process Specific Aspects
The performance of a membrane cell depends
on the following operating conditions:
1) Concentration of anolyte and catholyte
2) Current density
3) Temperature
4) Brine impurities
The optimum caustic strength depends on
the composition of the membrane polymer. To
achieve stable operation with high current efficiency, fluctuations in operating conditions or
upsets must be avoided. Fluctuations in the caustic strength beyond the optimum range influence
both the current efficiency and the cell voltage,
as shown in Figure 46 [131]. Dilution of the
anolyte caused by an upset in the brine feed also
decreases the current efficiency. The sensitivity
of membrane performance to operating conditions is attributed to changes in the water content
of the membrane.
CE (%,NaOH)
= 100−ηO2 −ηClO3 −ηClO+ηNaOH+ηNa2 CO3
where ηO2 , ηClO, and ηClO3 represent the
loss of current efficiency due to the generation of oxygen hypochlorite, and chlorate, while
ηNaOH and ηNa2 CO3 take into account NaOH
and Na2 CO3 introduced in the feed brine.
The current efficiency is mainly dependent
on membrane performance. On the cathode side
the membrane is in contact with a concentrated
NaOH solution, while the anode side has a pH of
Figure 46. Dependence of cell voltage and current efficiency on NaOH concentration [139]
Chlorine
7.2.1. Brine Purification
The introduction of membrane technology into
chlor-alkali electrolysis has dramatically increased the demands on brine purity [140]. The
lifetime of chlor-alkali membrane cells is determined by the operating conditions and the quality and purity of the feed into the electrolyzers.
Good long-term performance of the cells may
be obtained if brine impurities are kept within
the limits recommended in Table 14.
A major source of performance decline is
the accumulation of solid material in the membrane [141]. Specific impurity levels are dependent on membrane design, cell design, operating
conditions, the impurity itself and other impurities present. The prerequisite for long membrane life is to maintain low levels of, for example, Ca2+ , Mg2+ , Sr2+ , Ba2+ , Al3+ , SO2−
4
and SiO2 in the brine. Traces of these impurities damage the membrane and/or electrodes and
result in irrecoverable decreases in current efficiency and/or increased cell voltage. In the case
of a closed brine loop with no purge, each impurity brought into or formed in the system must be
removed to keep it below its specification level
and to prevent accumulation.
The contaminants can be brought into the
brine system by salt, by chemicals used in brine
purification steps, by water for dissolving the
salt, from materials of tanks, pipework, and cell
components, or by the process itself [141]. The
impurities in the salt depend upon the origin of
the raw material. Rock salt, vacuum salt, sea salt,
brine from well mining, or salt from waste incinerators serve as supplies of NaCl. The more
varied the sources are, the more diverse the impurities.
Membrane and electrode damage effect cell
performance, i.e., cause lower current efficiency,
increased cell voltage, and, as a result, increased
power consumption [142]. Some impurities affect the anode or cathode coating and cause an
increase in overvoltage or simply deposit in the
membrane, increasing its resistance and thus the
cell voltage. The increase in voltage may in some
cases be partially reversible when the impurity
concentration drops to the recommended limits.
Current efficiency declines are strictly related
to the membrane. Impurities lower the current
efficiency by reducing the membrane’s ability to
reject anions, specifically the ability to prevent
47
hydroxyl ions from migrating from the cathode
compartment through the membrane to the anode compartment [143]. This is usually a result of physical damage caused by precipitation and crystallization of impurities inside the
membrane. Impurities precipitate because the
environment in the membrane changes from an
acidic salt solution (pH 2 – 4) to a caustic solution (pH 14 – 15) over the 100 – 300 µm thickness of the membrane.
It is important to consider not only the impurities themselves but also their interaction. The
presence of one impurity may not be harmful,
but its synergistic combination with others may
be [143]. For example, silica itself is not harmful
for membranes. Only in the presence of calcium
and aluminum do precipitates form and damage
the membrane irreversibly. The concentration of
silica and/or the concentration of aluminum plus
calcium can be adjusted to give the optimum operating conditions. For example, with an effective secondary brine purification, higher levels
of silica can be tolerated. Similarly, if aluminum
concentration is high, calcium or silica concentration must be reduced to maintain acceptable
membrane performance.
To meet the strict requirements on brine purity outlined in Table 14 brine treatment is generally carried out in the following main steps in the
brine loop: saturation, precipitation, clarification, filtration, polishing filtration, ion exchange,
electrolysis, chlorate decomposition and dechlorination.
Calcium and magnesium are precipitated and
separated from the saturated brine with the insoluble materials. A ca. 10 wt % sodium carbonate
and barium carbonate (barium chloride) solution
and 32 wt % caustic soda are used as precipitants.
Ca2+ +Na2 CO3 → CaCO3 +2 Na+
2−
SO2−
4 +BaCO3 → BaSO4 +CO3
Mg2+ +2 NaOH → Mg(OH)2 +2 Na+
Alternatively:
−
SO2−
4 +BaCl2 → BaSO4 +2 Cl
2+
Hg
Al
2+
< 0.1 ppm
< 0.5 ppm
< 4 ppm
Ca +Mg
< 20 ppb
2+
parallel
< 0.2 ppm
operation
heavy metals
of an amalgam
plant
salt
salt
Ba2+
3+
salt
salt
Sr2+
Mg
2+
< 20 ppb
-
-
X
Ba2+ ;
SO2−
4
Al3+
SiO3− ;
Ca2+
SiO3− ;
Al3+
Hg2+
X
-
X
X
-
-
X
X
-
X/-
X
X
X
X
Ba2+ ; OH−
Ba2+ ; I−
Al3+ ;
SiO2−
3
X
-
-
X
X
-
-
caustic
X
X
brine
Solubility in
Sr2+ ; OH−
Mg ;
OH−
2+
Ca2+ ;
SO2−
3
Ca2+ ;
SO2−
3 ;
AL3+
Ca2+ ; I−
Ca2+ +Mg2+ Ca2+ ; OH−
Ca2+
salt
Impurity Source
Reagents
Max. limit
(w/w)
Table 14. Impurities and their effects
Damage
binding of
ion-exchange sites in
the membrane
physical disruption of
the membrane
-
formation of crystals
(zeolites, sodalites,
faujacites) near the
cathode side of the
membrane
precipitation near
cathode
side of the membrane
and crystalization
deposition on the
cathode
partially reversible,
covering of active
cathode coating
disruption of the
membrane
disruption of the
membrane
-
very fine precipitation minor damage on the
in the membrane
membrane, minor
interaction with
ion-exchange sites
coating of the anode
very fine precipitation
in the membrane
fine precipitation near
the anode side of the
membrane
precipitation on the
cathode side of the
membrane, formation
of crystals
precipitation near the physical disruption of
cathode
the membrane
side of the membrane,
formation
of large crystals
(blister formation)
Mechanism
+
++
+
+
++
Negative effect on performance
Voltage increase
An.
Cath.
Mem.
++
++
+
++
++
++
++
CE
PQu
precipitation with
Na2 S
Purge
precipitation as
hydroxide at pH 7 – 9,
ion exchange under
acid conditions
purge;
precipitation with
NaHSO3 plus ion
exchange
precipitation with
NaOH plus ion
exchange
coprecipitation with
Na2 CO3 plus ion
exchange
precipitation with
Na2 CO3 plus ion
exchange
Methods of control
48
Chlorine
X
∗ An. = Anode; Cath. = Cathode; Mem. = Membrane; CE = current efficiency; PQu = product quality.
evolution of N2
< 0.1 ppm
Fe3+
K4 (Fe(CN)6 ) anticaking-agent
for salt
TOC
increased foaming,
overplating
as Fe3+
-
destruction of
ion-exchaange resin
precipitation on the
physical disruption of
cathode side of the
the membrane
membrane, formation
of crystals
very fine precipitation
in the membrane
destruction of the
anode coating
precipitation near the reduction of the OH−
cathode surface of the ion rejection
capability
membrane
coating of the anode
formation of CO2
< 1 ppm
Fe3+
+
SO2−
4 ; Ba
CO2−
3 ; H2 O
X
Damage
deposition on the
covering of active
cathode (in extreme coating (puncturing
cases: dendritic
of the membrane)
growth from cathode
toward the anode)
absorption of Ni in the
membrane, deposition
on the cathode
Mechanism
salt
< 0.4 g/l
Na2 CO3
Na
+
X
-
-
X
X
caustic
chlorination of
ion-exchange resin
salt,
precipitation
with Na2 CO3
or BaCO3
process, side
reactions
salt,
< 4 – 8 g/l
dechlorination Na2 SO4
with NaHSO3
SO2−
4 ;
X
X
-
I− ; Ca2+
I− ; K +
X
X
X
brine
I− ; Na+
Ni2+ ; OH−
Fe3+
Reagents
Solubility in
< 10 g/l
NaClO3
ClO−
3
CO2−
3
SO2−
4
F−
< 0.5 ppm
salta,
< 0.2 ppm
pipework, tank heavy metals
material,
cathode
salt
< 0.2 ppm
Ni2+
salt
salt, pipework, < 0.1 ppm
tank material,
anti-caking
agent
Fe3+
I−
Source
Max. limit
(w/w)
Impurity
Table 14. Continued
+
+
++
++
+
++
++
+
+
Negative effect on performance
Voltage increase
An.
Cath.
Mem.
++
++
CE
++
++
PQu
oxidation with active
chlorine plus
precipitation with
NaOH
chlorate
decomposition by
acidification
filtration
purge, precipitation
with BaCO3 or BaCl2
plus ion exchange
purge
purge
ion-exchange, purge
precipitation with
NaOH
Methods of control
Chlorine
49
50
Chlorine
From the precipitation tank, the brine is fed
into the clarifier, where a defined quantity of
flocculant is added to promote the settling of the
precipitated solids and gels. The brine is then
pumped to a filtration system followed by an
ion exchange purification.
Additionally, if the brine circulating system
of an existing mercury cell plant also serves
membrane cells, all mercury must be removed
in a chemical treatment facility. The brine from
the primary filtration is acidified with hydrochloric acid to pH 2.0 – 2.5 and sodium sulfide is
added to precipitate mercury sulfide. Subsequently, the brine is filtered, alkalized to pH 9.5 –
11 by adding caustic soda, and finally fed to the
secondary purification section.
The content of calcium and magnesium must
not exceed 20 ppb. Such low contents can be
achieved by using ion-exchange columns. The
polished brine is pumped at approximately
70 ◦ C to an ion-exchange system with two resins
beds operating in series according to the lead/lag
principle. When the leading ion exchanger is exhausted it is put to the regeneration and conditioning mode, while the lagging one takes over
the lead position. After treatment in this secondary purification step, the purity limits are
met and the brine is fed to the membrane cells.
In another secondary purification system two
columns operate in series while the third is in
regeneration mode. When the first column is exhausted, the regenerated column is put in second
position.
Prior to the resaturation, the byproduct chlorate and dissolved chlorine must be eliminated
from the anolyte. Chlorate concentration is controlled by acidification of a partial stream of
anolyte with an excess of hydrochloric acid
[143]. Chlorine is removed under vacuum followed by addition of sodium bisulfite and hydroxide.
NaClO3 +6 HCl → NaCl+3 Cl2 +3 H2 O
2 Cl2 +2 NaHSO3 +6NaOH → 4NaCl+4 H2 O+2 Na2 SO4
All other impurities not precipitated, filtered
out, or extracted by the ion exchangers can only
be controlled by purging a partial stream of
the anolyte to avoid accumulation. Resaturation
then closes the loop.
7.2.2. Commercial Membranes
The ion-exchange membrane is the key component of the membrane cell. The energy consumption and the quality of the products depend on
membrane performance. Requirements for the
membrane are as follows:
1) Durability under the conditions of chlor-alkali
electrolysis
2) High selectivity for sodium ion transport
3) Low electrical resistance
4) Sufficient mechanical strength for practical
use
5) Low sensitivity to changing operating conditions
The importance of 1 – 3 is described in Section 7.1. High mechanical strength is necessary
for installing the membrane and during service
life, in which the membrane has to cope with deviations in temperature, concentration, and pressure.
As the performance of the membrane is the
most important element in the economy of a
membrane cell, many refinements have been
made in membrane manufacturing. To reduce
the current screening due to fabrics, membranes
reinforced with dispersed microfibers and interwoven fabrics made of electrolyte-soluble fibers
and PTFE have been developed [144, 145].
The improvement of hydrophilicity by covering the surface on the cathode side or on both
sides with a nonconductive inorganic material
brought about a significant reduction in the cell
voltage. The surfaces of the membrane are covered with thin layers of a porous inorganic material. This material is an oxide, hydroxide, or
carbide of the metals of groups 4, 5 and 6 or the
iron triad (Fe, Co, Ni) [138, 146].
Figure 47 illustrates the effect of hydrophilic
cathode surface modification. The surfacemodified membrane (Type B) has a lower cell
voltage than the conventional membrane (Type
A). Tthe voltage of the surface-modified membrane decreases linearly with decreasing gap
size. With these advanced membranes, so-called
zero-gap cells have been made possible, and the
ohmic loss in electrolytes has been reduced to a
minimum.
Chlorine
The active life of a membrane is determined
by the economic balance between membrane
cost and energy cost in use [148].
Figure 47. Effect of cathodic surface modification
The performance of membranes depends on
the operating conditions, especially on the caustic strength of the solution (Figure 46). Commercially available membranes, delivered by Asahi
Chemical, Asahi Glass and Du Pont, are designated for use in a specific strength of caustic. For
economic production the selection of the appropriate membrane is essential. Table 15 gives an
overview of the most widely used membranes.
Most membranes are operated in the narrow- or
zero-gap configuration to minimize power consumption.
7.2.3. Power Consumption
For monitoring cell performance and comparing
different electrolyzer designs, the electric power
required to produce one tonne of NaOH 100 %
is considered. This figure is determined by the
voltage drop over one cell and the NaOH current
efficiency.
DC Energy
= UMIt
100% NaOH produced
kWh
= t
=
U
F ·CE
51
where U is the cell voltage (V), F the Faraday
constant for NaOH (1.4923 kg/kAh), and CE the
NaOH current efficiency (%).
The specific power consumption is the main
indicator for economic plant operation, and continuous efforts are made to lower the voltage
and increase the current efficiency. At a thermodynamic minimum the decomposition voltage
of about 2.2 V limits the theoretical minimum
energy requirement to about 1480 kWh/t 100
% NaOH, as shown in Figure 48. At practical
current densities of 3.0 – 5.5 kA/m2 for presentday commercial cells and membranes, power
consumption measured at the electrolyzer terminals is in the range of 1950 to 2180 kWh/t
100 % NaOH dependent on the selected current density (anolyte/catholyte temperature 90
◦
C, NaOH concentration 32 wt %, NaCl concentration in the anolyte 220 g/L). The power consumption rises with increasing operating time
due to aging effects, such as decreasing current efficiency and increasing voltage. Investment costs rise when operating at low current
densities, as more cells are needed to meet production. Hence electrolyzers are operated at low
current densities in countries with high energy
prices, and at high current densities in countries
with low energy prices.
7.2.4. Product Quality
The caustic soda solution has a concentration of
up to 32 ± 1 wt % NaOH. If a NaOH concentration of 50 wt % is required, evaporation can be
used. The typical NaCl content is 20 ppm in a
32 wt % caustic solution.
The hydrogen has almost synthesis quality
with a concentration of about 99.9 vol % H2 (dry
basis). The chlorine has an oxygen content of
about 1.5 vol % (dry basis). Chlorine with an
oxygen content below 0.6 vol % (dry basis) can
be obtained by acidifying the brine with hydrochloric acid.
7.3. Membrane Cells
7.3.1. Monopolar and Bipolar Designs
A commercial membrane plant has multiple cell
elements combined into a single unit, called the
52
Chlorine
Table 15. Commercial membranes
Asahi Chem F5201
Aciplex
F4202
F4203
F890
Asahi Glass F892/old
Flemion
F892/new
F893/new
N90209
DuPont
N966
Nafion
N981
a
Available
since
Tear strength,
Tensile strength,
kg
kg/cm
1991
1993
1997
1989
1990
1994
1994
1984
1988
1996
4
4
4
4.5
4.5
4.5
4.5
2.5
5.5
5.5
5.5
5.5
6
5
5
5
5.6
7.3
3.3
Ohmic drop
(at 3 kA/m2 , T = 90 ◦ C,
c = 32 wt %),
V
Caustic strength,
a
33 – 36
30 – 34
30 – 34
31.5 – 32.5
30 – 35
31.5 – 33.5
31.0 – 32.5
30 – 35
30 – 35
30 – 35
a
a
0.35
0.27
0.28
0.26
0.35
a
a
wt %
Not published.
Figure 48. Specific power consumption
electrolyzer. The electrolyzers follow two basic
designs: monopolar and bipolar [147].
In a bipolar arrangement the elements are
connected in series with resultant low current
and high voltage. The cathode of a cell is connected directly to the anode of the adjacent cell,
as shown in Figure 49. The operation of a bipolar electrolyzer can be easily monitored by measurement of element voltages. If element upsets occur, a safety interlock system actuates
the breakers (short-circuiting switches) and isolates the electrolyzer from the electric circuit. As
the influx and efflux of electrolytes for the cells
with different electric potential are gathered in
common headers, problems of stray current may
arise.
In the monopolar type all anodes and cathodes are connected in parallel, forming an electrolyzer with high current and low voltage (Figure 50). Due to the long current path, the voltage
drop is high and can only be reduced by minimizing the size of cells or introducing internal
copper conductors to lower the resistance. Because of this basic principle, ohmic losses in the
monopolar cells are 80 – 100 kWh per tonne 100
% NaOH, which is much higher than in equivalent bipolar cells. Furthermore, the bipolar safety
system is not applicable to the monopolar design, since the cell elements are arranged in parallel, which does not permit the monitoring of
deviations in individual cell voltages.
Figure 49. Bipolar electrolyzer
Chlorine
Figure 50. Monopolar electrolyzer
Multiple electrolyzers are employed in a single d.c. circuit (Fig. 51). Usually bipolar electrolyzers are connected in parallel with low current and high voltage. Monopolar electrolyzers
are often connected in series, resulting in a high
current circuit and low voltage. Though both
principles still appear on the market, investment
and operating cost considerations, such as for the
rectifier system, the cell room space required,
for piping, valves, instrumentation, busbars and
switches, significantly favor the bipolar design.
53
ranges from 0.2 to 5.0 m2 . Current density varies
between 1.5 and 7 kA/m2 .
The cells are filled with electrolytes, and gasseparating means are provided outside the cells.
Many cells generally stacked like a filter press,
constitute one electrolyzer with high production
capacity.
The performance of a plant is determined by
the electrolyzer, the cell voltages, and the current efficiency of the membrane. It is essential
to design an electrolyzer with an homogeneous
electrolyte concentration, temperature, and current density distribution across the whole area
of the membrane.
The construction materials of the cell are selected to withstand the corrosive electrolytes. In
most electrolyzers, titanium and nickel are used
for the anode and cathode compartments of the
cell. In older electrolyzers, stainless steel is used
on the cathode side.
For economic and environmental reasons,
mercury and diaphragm plants are increasingly
being converted to membrane electrolyzers. The
existing facilities, such as rectifiers, equipment
for brine purification, and equipment for product
treatment are utilized as much as possible.
7.3.2. Commercial Electrolyzers
Asahi Kasei ACILYZER-ML/NC Electrolyzer. The Asahi bipolar electrolyzer
(Fig. 52) is of the filter-press type. The bipolar cell frames are suspended in a steel frame
and compressed by a hydraulic device. Each
cell frame consists of an anode and cathode
compartment separated by a partition wall. The
anode compartment is made of titanium, and
the cathode compartment consists of special
stainless steel and nickel. The anode and cathode structures are spot welded onto ribs in each
compartment. Each compartment has an inlet
nozzle for electrolytes at the bottom and an
outlet nozzle for gas and electrolyte on top,
connected to the gas/liquid separation chamber.
Two types of cell frames are available: frames
with forced circulation of electrolytes by pumps,
and frames with natural circulation in each compartment by means of a special arrangement of
integrated ducts. The current is connected to the
first and last element by flexible busbars.
Generally, membranes are clamped vertically
between the meshlike metal anodes and cathodes. The effective membrane area of a cell
CEC BITAC 800 Electrolyzer. The Chemical Engineers Corporation (CEC) bipolar
BITAC electrolyzer was jointly developed with
Figure 51. Electrolyzer architecture
54
Chlorine
Figure 52. Cell structure of ACILYZER ML32NC
a) Gasket; b) Nickel; c) Cathode; d) Anode; e) Titanium; f) Partition wall; g) Membrane; h) Rib; i) Reinforcing rib; j) Duct;
k) Gas – liquid separation chamber
Tosoh Corporation. The design follows the filterpress principle. Up to 80 bipolar electrode
frames are clamped together by end plates and
spring-loaded tie rods. The frames are made of
special titanium alloy for the anode and nickel
for the cathode. The electric current flows along
the nickel pans, since the electrical conductivity
of nickel is six times higher than that of titanium.
Gas and electrolytes leave the cell compartment
in overflow mode with little pressure fluctuation.
Transparent PTFE tubes are attached at the electrolyte inlet and outlet nozzles of each element.
Anolyte recirculation takes place through an external loop.
CEC CME DCM 400 Electrolyzer. The
CME monopolar electrolyzer consists of large
elements compressed in a filter-press arrangement. The electric current travels into each anode element through conductor rods and current
distributors. This design achieves uniform current distribution over the large electrode area.
The current distributors serve the additional
role of a downcomer pipe, which creates a natural circulation within the cell, providing a uniformly distributed electrolyte concentration as
well as good gas release. The anode frames
are constructed from titanium and the cathode
frames from a special stainless steel. The rods
are cladded with titanium and stainless steel.
Inlet and outlet tubes for liquids and gases and
are made of transparent PTFE.
Uhde BM 2.7 Electrolyzer. The
bipolar
Uhde electrolyzer (Fig. 53) is a single-element
concept. Each element comprises anode and
cathode half-shells, electrodes, a membrane,
flanges, and the sealing system. This enables
long-term storage of pre-assembled and fully
tested elements. The electrodes are attached
with continuous laser weld to the current transfer and support blades and hence to the halfshells. The anode is made of titanium and the
cathode of nickel. The individual cell elements
of an electrolyzer are suspended in a steel frame
in which they are lightly pressed together for
electrical contact. Large sealing forces are not
required in the single-element concept, as each
element is a separate, stand-alone electrolysis
cell. The feed and discharge lines of the cell are
located underneath the cells and connected to the
catholyte and anolyte headers. The area above
the electrolyzer is free of piping or bus bars,
simplifying access and eliminating the risk of
leakage and associated corrosion problems. The
current is conducted from cell to cell by continuously laser-welded, explosion-bonded titanium
– nickel contact strips on the anode half-shell.
Chlorine
55
Figure 53. Uhde BM 2.7 electrolyzer
a) Single element; b) Contact strip; c) Cell rack; d) Busbars; e) Inlet hoses; f) Outlet hoses; g) Header
The brine and caustic soda feeds enter the cell
at the bottom, and the product streams are discharged downwards through internal overflow
pipes. The internal baffle plate at the top of the
anode half-shell prevents gas-phase blistering
of the membrane. The chlorine gas is effectively
removed from the membrane, preventing contact and improving the inherent safety of the
electrolyzer. Natural circulation around a downcomer plate and a distribution pipe for brine and
caustic achieve homogeneous temperature and
concentration profiles within the element and
assist in achieving uniform current distribution.
EL-Tech ExLB Bipolar Electrolyzer. The
ExLB bipolar electrolyzer (Fig. 54) is basically
of the same design as the EL-Tech ExLM electrolyzer. Instead of the copper distributors with
interface material, which provide the parallel
(monopolar) arrangement in the electric circuit, the elements are connected in series (bipo-
lar), omitting the copper distributors and simply
pressing the nickel cathode pan onto the nickelplated back of the anode pan. The integral feed
and discharge manifolds are designed to avoid
current leakage.
EL-Tech ExLM Monopolar Electrolyzer.
The ExLM monopolar electrolyzer is an improved version of the MGC electrolyzer, which
has been in service for more than 15 years.
The elements are sealed with O-rings in a staggered gasket design. The cathode O-ring is located closer to the liquid than the anode O-ring.
This protects the anode O-ring from the chlorination degradation, making it a long-life back
up seal. The elements are pressed together by
tie rods with copper distributor plates and conductive interface material to provide good current distribution. Electrolytes and gases are fed
and discharged to and from the elements through
the manifold passage attached to the cell ele-
56
Chlorine
Figure 54. EL-Tech ExLB bipolar electrolyzer
ments. Increased internal electrolyte circulation
is achieved by an improved electrode design.
EL-Tech ExLDP Dense Pak Unit Electrolyzer. The EL-Tech ExLDP dense pak electrolyzer comprises monopolar sections in one
electrolyzer filter press compression set using standard monopolar cell components. Each
monopolar cell section is separated by an insulating Inter Pak Spacer. Mostly three monopolar
electrolyzer sections are included, with 2 to 10
elements per section. The dense pak can be configurated to match special rectifier/transformer
configurations of existing plants, making it suitable for mercury and diaphragm cell conversion projects. For new plants, the advantage of
the ExLDP electrolyzer is reduced current, increased voltage circuits compared to an equivalent monopolar cell unit.
INEUS FM21-SP Electrolyzer. The FM21SP (Fig. 55) is a monopolar electrolyzer incorporating a simple pressed electrode structure. The
anode assembly is composed of a 2 mm thick
titanium panel between compression molded
joints of a special cross-linked EPDM elastomer.
The cathode assembly is composed of a 2 mm
thick nickel panel between compression molded
joints, also of EPDM.
The anodes and cathodes are assembled between 2 end plates until the number of electrodes required for the desired electrolyzer capacity is reached, up to 60 anodes in the FM21SP and up to 90 anodes in the larger FM1500.
A key feature of both designs is the elimination
of any external piping to individual cell compartments by the use of a simple but effective
internal header/manifold arrangement.
Chlorine
57
Figure 55. INEUS FM21-SP elelctrolyzer
a) Tie-rod; b) Floating end plate; c) Copper electrical connections; d) Ion exchange membrane; e) Fixed end plate; f) Anode
electrode assembly (titanium panel between compression molded gaskets; g) Cathode electrode assembly (nickel panel between compression molded gaskets); h) Support rail
The electrolyzer has coated titanium anodes.
The cathodes are pure nickel, also available with
a coating to lower the hydrogen overpotential if
necessary. Both electrodes are pressed from integral sheets of pure metal, and this makes recoating of the electrodes extremly simple and cost
effective. Hence recoated structures can be sent
to site prior to electrolyzer refurbishment from
a pool of electrodes available to all customers.
Effective electrode area is 2 × 0.21 m2 per
electrode, which gives a very compact electrolyzer. The individual electrodes are readily
handled without the need for lifting apparatus,
which allows the electrolyzer to be rebuilt and
refurbished in the minimum of time.
The media are fed to and discharged from the
electrolyzers by a header system arranged along
the walls of the cell room on one side.
From the other side power is supplied either
from separate transformer/rectifier units for each
electrolyzer or from one unit for two or more (up
to six) electrolyzers in parallel. The switches are
arranged close to the rectifiers. They are actuated
automatically and connected to the common interlock system for safety reasons.
In the middle space remains available for the
electrolyzers and their individual feed and discharge piping. Only a light crane is required to
handle single electrode frames or elements. Thus
only a light structure for the entire cell house is
used.
7.3.3. Comparison of Electrolyzers
8. Electrodes
Operating parameters of bipolar electrolyzers
are compared in Table 16, and those of monopolar electrolyzers, in Table 17.
7.3.4. Cell Room
Typical bipolar membrane cell rooms are shown
in the following Figures 56 and 57.
8.1. Anodes
The initial anodes used for the electrolytic generation of chlorine were made of platinum or
magnetite. However, as the plant grew in the size,
the cost of platinum and limitations of the current
density for magnetite led to the wide-scale introduction of graphite anodes, which were used exclusively up to 1970. The graphite of choice was
58
Chlorine
Table 16. Bipolar electrolyzers
Company
Cell
Effective membrane
area, m2
Max. no. of elements
Current density,
kA/m2
Max. capacity of
electrolyzer t/d NaOH
100 %
d.c. power
consumption kWh/t
NaOH
(at current density)
CEC
KRUPP UHDE
EL-TECH
ML 32
ASAHI KASEI
ML 60
BITAC 800
BM 2.7
ExLB
2.72
5.05
3.276
2.72
1.5
150
up to 6.0
150
up to 6.0
80
1.5 – 6.0
160
1.5 – 6.0
80
1.5 – 7.0
45
90
54
90
29
2100
2100
2150
2130
2100
(4.0)
(4.0)
(5.0)
(5.0)
(5.0)
Table 17. Monopolar electrolyzers
Company
Cell
2
Effective membrane area, m
Max. no. of elements
Current density, kA/m2
Max. capacity of electrolyzer, t/d
NaOH 100 %
d.c. power consumption kWh/t
NaOH
(at current density)
CEC
INEUS
EL-TECH
CME DCM 400
FM 21-SP
ExLM
3.03
32
1.5 – 4.0
13
0.21
120
1.5 – 4.0
7
1.5
30
1.5 – 6.0
9
2150
2140
2150
(3.5)
(4.0)
(5.0)
Figure 56. Bipolar cell room by Ashai Glass
Chlorine
59
Figure 57. Bipolar cell room by Krupp Uhde
low in ash and vanadium and composed of various types of particulate coke and pitch binder.
Following extrusion, baking at ca. 1000 ◦ C, and
graphitization at 2600 – 2800 ◦ C, the final shape
of the electrode was achieved by machining. The
shape of the horizontally suspended anodes with
an initial thickness of 7 – 12 cm for the amalgam
process was similar to that of modern titanium
anodes due to the retrofitting of existing cells.
The anodes had vertical slits and holes to allow
the removal of the gaseous chlorine. Due to the
cogeneration of oxygen and the resulting formation of CO and CO2 , electrode wear was high,
in the range of 1.8 – 2.0 kg graphite per tonne
of chlorine from NaCl and 3 – 4 kg per tonne
from KCl. Even with a daily adjustment of the
anodes to compensate for the changes in dimension a k value of only 0.12 to 0.14 Vm2 kA−1
was achievable.
The initial attemps to replace the graphite
anodes with activated titanium anodes began
as early as 1957 with platinized titanium and
Pt/lr-coated anodes. However because of the
short lifetimes of the anodes, they were not economic. The use of mixed metal oxides was first
patented by Beer in 1965 and 1967 [149]. The
initial patent described a coated metal electrode
in which the active material was a mixed metal
oxide coating containing one or more of the platinum metal group oxides. The second patent described coatings in which mixed metal oxide
crystals contained a non-platinum metal oxide
in addition to the platinum metal oxide (including Ti, Ta, and Zr oxides).
Further improvements in the coating and the
anode structures followed rapidly along with the
commercialization of anodes by De Nora [150]
under the trade name Dimensionally Stable Anode (DSA). Because of the dimensional stability
and the lifetime of the coating and the ability to
increase the current densities, rapid introduction
of the activated titanium anodes was possible. At
present only a few plants still use graphite anodes, largely due to the initial investment costs
for titanium anodes.
8.1.1. General Properties of the Anodes
Coating Properties and Preparation.
Comprehensive reviews on preparation and
properties are given in [151, 152].
Chemical Composition. Because of its
price and performance, Ru is the basic component in all commercial coatings at present,
60
Chlorine
along with an oxide of a non-platinum metal
(e.g., Ti, Sn, or Zr). In most cases a second
platinum metal oxide is added to increase the
performance of the anode coating. There is an
optimum ratio of platinum metal oxides to nonplatinum metal oxides in terms of overpotential,
wear rate, and costs. The optimum depends
on the operating conditions and the method of
preparation of the coating and normally lies in
the range of 20 : 80 to 55 : 45 by weight. Some of
these coatings may contain glassy fibers [153]
and some contain pre-oxidized material such as
Li0.5 Pt3 O4 [154].
Preparation. The solvent used for the preparation of the precursors solutions are chosen on
the basis of the desired electrochemical properties and the method of application, which is
mainly determined by the anode structure. Most
coating solutions are prepared by dissolving
salts or organometallic complexes in aqueous,
organic, or mixed solvents. The coating can be
applied by spraying, brushing, dipping, or other
techniques. Following evaporation of the bulk of
the solvent, the anode is heated to 350 – 600 ◦ C
to form the oxidic coating then cooled prior to
the next coating cycle. This is repeated until the
desired coating thickness is applied. Optional
post-thermal treatment can also be carried out.
The optimal performance of the coating depends on the above parameters and on the coating thickness per coating cycle, which must be
optimized for each coating and surface pretreatment step.
Crystallographic Composition, Morphology, and Real Surface Area. A rutile phase is
the electrochemically active phase of the coating, and although it is thermodynamically unstable, it remains even after many years of operation. The stable phase – anatase TiO2 in the case
of TiO2 – RuO2 coatings – is electrochemically
inactive [155]. The degree of crystallinity and
the composition are related to the processing parameters [156] and the various degrees of mixed
crystals exhibit different stabilities. The real surface area of the coating is a function of both
the titanium pretreatment and the coating composition. The surface of the chlorine-generating
coating is often described as “cracked mud” due
to its resemblance to a dry river bed. The BET
surface area of the coatings or that determined
electrochemically vary ca. 400 to 1000 times the
geometric surface area [157].
Overpotential and Current – Voltage Relationship. The observed overpotential for chlorine evolution at 2 – 10 kA/m2 is in the range of
80 – 110 mV [158 – 161], about 70 – 100 mV of
which is due to diffusion overpotential effects
[161]. The overpotential for the generation of
oxygen under similar pH and temperature conditions lies is ca. 300 mV more anodic than that
of chlorine generation. Other than oxygen evolution, the only other side reaction is formation
of chlorate.
Coating Wear and Coating Lifetime. The
coating lifetime is strongly dependent not only
on the type of cell – membrane, diaphragm, or
mercury – but also on a range of process parameters, including brine quality, current density, and
membrane or diaphragm quality. The upper limit
of the wear rate would seem to be in the region
of 500 t Cl2 /m2 anodic area for a standard commercial loading.
The wear rate mechanism is discussed in detail in [162, 163]. The effects of various impurities and materials in the brine can be divided
into three types.
1) Compounds or ions which attack the substrate, e.g., fluoride or organic acids such as
formic or oxalic acid.
2) Materials which built up blocking layers on
the surface of the anode, e.g., hydraulic oil
or polymer films resulting from delamination
of membranes. The irreversible poisoning of
coating is caused by ultrathin aluminum silicate layers.
3) Electrochemically active film-forming materials such as MnO2 , which may lead to an
increased oxygen content in the chlorine.
Other examples, such as the insensitivity of
the performance of diaphragm anodes to almost
complete surface coverage by iron oxides illustrate the robustness of commercial coatings.
8.1.2. Anodes for Mercury Cells
Structure. The classical structure of anodes
for this process still reflects the retrofitting concept used during the 1970s and the high current
Chlorine
operations at ca. 10 kA/m2 . A typical mercury
cell anode consists of a number of copper shafts,
protected by either a permanently welded or removable titanium outer sleave, from which the
current is distributed to the active surface over
distributor bars (Fig. 58).
61
value < 0.5 mm is achieved, mostly by manual
straightening after manufacturing or recoating.
Coating Life. The coating life is determined
by a wide range of practical aspects and normally not directly related to the electrochemical
wear rate of the coating. These include:
– Mechanical damage to the anode caused by
short circuiting [165, 166]
– The need to maintain a recoating schedule,
due to production demands and the labor intensive refitting of a cell.
– Synchronization of recoating with the exchange of other consumable parts of the cells
such as covers and gaskets.
Figure 58. Four-stem anode for amalgam cells
a) Active surface; b) Current distributor; c) Riser tube to
protect the copper bar inside
The quick release of gas and the supply of
fresh brine to the active surface are the major
requirements of an mercury cell anode, and a
wide range of designs have been built. The most
common types are shown in Figure 59. The differences are more evident at current densities
> 7 kA/m2 . The use of baffles on the back of
the active surface to enhance the gas lift and aid
the supply of brine to the active surface is also
common [164].
Flatness is critical for the optimal performance of the anodes in the cells. A typical
The tendency has been to increase the lifetime
from about 180 t Cl2 /m2 to 300 – 400 t Cl2 /m2 .
This has been achieved by the introduction of
better control systems in the cells and the development of intermediate layers of plasmasprayed conductive TiO2−x between the active
coating and the titanium substrate [159].
8.1.3. Anodes for Diaphragm Cells
The predominate determinants in the design of
diaphragm anodes are:
– The relatively low current density of ca.
2.0 kA/m2
– The minimization of the anode – diaphragm
gap
– The need to remove and replace the anode
array from the cathode, hence the use of retractible anodes
Figure 59. Anode designs for quick gas release
A) Flatt profile (channel blades); B) Rod type (3-, 4-, 5-mm diameter); C) 3D side profile anode
62
Chlorine
– The limiting height of the cathode, integrated
with the diaphragm manufacturing technology
The conventional anodes shown in Figure 60
[167] have been further developed to optimize
the energy consumption of the cells by replacing the simple flat expanded metal with complex
structures [169]. At present very few plants are
still operating without expandable anodes.
Figure 61. Empirical fit of observed nonlinear wear rate of
coating thickness versus years on line.
L(t) is the loading at time t, L the initial loading, r the wear
rate, t time, A(t) the active surface area at time t, q an empirical factor related to the current density sensitivity of the
wear rate r
Figure 60. Anode for monopolar diaphragm cells
a) Activated (coated) expanded metal; b) Expanding spring;
c) Titanium-clad copper bar; d) Copper thread to fix the anode to the cell base
Another type of diaphragm anode is used in
the bipolar Ganor cells [168].
Coating Life and Mechanism of Deactivation. The coating lifetime of DSA coatings exceeds 12 years, and production of chlorine exceeds 240 t Cl2 /m2 . The wear is caused by the
relatively high oxygen content in a diaphragm
cell [170] of ca 1 – 2 %. The wear rate is nonlinear (Fig. 61). This nonlinearity is critical for
determining the correct time to begin recoating
so as to avoid unplanned stoppages.
8.1.4. Anodes for Membrane Cells
Structure. The variety of designs of membrane cells has led to a range of anodes active
area structures; the common principles are the
need to support the membrane and gas release
to the back of the anode surface. Therefore, thin
flattened expanded, perforated metals or louver
type structures with and without perforations are
used [171].
Coating Life. At present the second-generation coatings for membrane cells are showing
lifetimes comparable to those of the diaphragm
process. The actual lifetime of the anodes is dependent on the extent of damage by caustic flow
through holes in the membrane or by contamination with the poisons.
Oxygen Content [151]. The oxygen content
has also been improved by the optimization of
internal circulation of the brine within the cells.
8.2. Activated Cathode Coatings
Since 1910 diaphragm brine electrolyzers have
used carbon steel cathodes and continue to use
carbon steel to this day. When the first ionexchange membrane electrolyzers were introduced in the late 1970s, the cathodes were also
carbon steel. By the early 1980s the design
had evolved to stainless steel and nickel cathodes, and finally in the 1990s to exclusively
nickel cathodes. Depending on current density,
the hydrogen overpotential of carbon steel cathodes is about 300 mV. Active cathode coatings
can lower the overpotential by 200 – 280 mV,
thus providing significant energy savings. Active coatings have often been described in the
literature and used in water electrolysis for over
40 years. With the development and evolution of
the ion-exchange membrane technology, active
cathode coatings are coming into general use.
Chlorine
The patent literature covers many different
types of coatings, and new ones are being published regularly. The two basic approaches to activation are high-surface area coatings and catalytic coatings. Both bare nickel and carbon steel
show lower hydrogen overpotential once in operation and their surfaces roughen. In fact by
grit blasting bare nickel cathodes and roughening the surface, the long-term overpotential can
be reduced by 30 – 40 mV. More common are
porous nickel-type coatings that offer high surface area and good chemical resistance. These
coatings consist of two or more components.
At least one of the components is leached out
in caustic to leave the porous high surface area
nickel [172]. These coatings are typicall nickelzinc [173], nickel – aluminium – Raney nickel
[174], nickel – aluminium [175], or nickel – sulfur [176]. A variety of additives are recommended for strength, life, and resistance to poisoning by impurities. Rough coatings of nickel –
nickel oxide mixtures [177] and nickel with embedded activating elements such as ruthenium
[178] are also used. Sintered nickel coatings
are described in patents [179] as well as being
available from Huntington Alloys. Nickel coatings containing platinum group metals, primarily platinum and/or ruthenium, have been sold
by Dow [180], Johnson Matthey, and ICI [181].
The coatings used for diaphragm and membrane electrolyzers differ because of the different substrates (carbon steel and nickel, respectively) and the different operating conditions.
The weak 11 % caustic in diaphragm cell liquor
is less corrosive than the strong 33 % caustic
of a membrane electrolyzer. The less expensive
and more fragile coatings like nickel – zinc can
be used in diaphragm electrolyzers. Membrane
electrolyzer suppliers favor the platinum group
metal coatings.
The shape of the cathode structure is an important factor affecting the choice of cathode
coating. The complex cathodes of diaphragm
electrolyzers lend themselves to liquid systems (e.g., electroplating or electroless baths)
that can coat the entire structure by immersion
[182]. Membrane electrolyzers, which are primarily of a filter-press design, have flat cathodes that are easy to coat by spraying or painting. The lower operating current density of diaphragm electrolyzers means more cathode area
per unit of production; this requires a less ex-
63
pensive coating. Most diaphragm electrolyzers
use heat-cured polmer – asbestos separators (diaphragms) that are vacuum deposited after the
cathode coating is applied. This curing operation can destroy the activity of certain coatings.
All cathode coatings are susceptible to poisoning by impurities that make their way into the
catholyte with the deionized water or are components of the piping, electrolyzer etc. These impurities tend to blind the activity of the coatings
over a period of time that depends on their concentration. Porous nickel coatings in diaphragm
electrolyzers are less susceptible to blinding by
impurities because spalling of the brittle coating
makes the coatings self-cleaning. The platinum
group metal coatings are subject to damage from
reverse currents during electrolzer outages. Precautions are needed to protect the coatings with
reducing agents [183] or by cathodic protection
[184].
Active cathode coating have become the standard throughout the chlor-alkali industry for new
construction with the ion-exchange membrane
electrolyzer technology. In most of the older diaphragm electrolyzer plants, problems with application of the cathode coatings and generally
lower power costs have obviated the use of active
cathode coatings. While there are more recent
developments applicable to the diaphragm technology in the way of active cathode coatings,
many of these developments remain the proprietary information of the technology and coating
suppliers.
9. Comparison of the Processes
The advantages and disadvantages of the three
chlor-alkali processes are summarized in Table
18. The three chlor-alkali processes can be compared in respect to the quality of the chlorine and
caustic produced, and the equipment and operating costs.
Today the membrane process is the state of
the art for producing chlorine and sodium hydroxide or potassium hydroxide. All new plants
are using this technology. The production capacity of chlor-alkali plants using the membrane
process reached about 21 % of total world production capacity in 1995 and is predicted to increase to about 28 % by 2001 (Table 19) [132].
64
Chlorine
Table 18. Advantages and disadvantages of the three chlor-alkali processes
Process
Advantages
Disadvantages
Diaphragm process use of well brine, low electrical energy consumption use of asbestos, high steam consumption for caustic
concentration in expensive multistage evaporators, low purity
caustic, low chlorine quality, cell sensitivity to pressure variations
Mercury process
50 % caustic direct from cell, high purity chlorine anduse of mercury, use of solid salt, expensive cell operation, costly
hydrogen, simple brine purification
environmental protection, large floor space
Membrane process low total energy consumption, low capital investment,use of solid salt, high purity brine, high oxygen content in
inexpensive cell operation, high-purity caustic,
chlorine, high cost of membranes
insensitivity to cell load variations and shutdowns,
further improvements expected
The diaphragm cell capacity remains constant
and there is a decline in mercury cell capacity.
Table 19. World chlorine market 1995 and 2001 (in %)
Diaphragm
Mercury cell
Membrane
Others
Market, 106 t/a
1995
2001
52
22
21
5
41
49
18
28
5
43.3
The conditions for a conversion from the mercury and the diaphragm process to the membrane
process are discussed below.
9.1. Product Quality
Table 20 shows typical composition values for
the chlorine and caustic produced by the diaphragm, mercury, and membrane processes.
Chlorine produced by the mercury process can
be used directly for most uses. Chlorine produced by the diaphragm or membrane process
contains up to 2 % O2 , depending on the pH of
the anolyte. This oxygen can be removed by condensation and evaporation of the chlorine.
The sodium hydroxide solution from the mercury process is the purest of the three; the
amounts of NaCl and NaClO3 are especially low.
However, the quality of caustic from the membrane process is almost as good. A main drawback of the diaphragm process is the high concentration of NaCl and NaClO3 in the caustic
solution. This sodium hydroxide solution cannot be used for some processes. A chloridefree grade, commonly referred to as rayon-grade
caustic, is required for 20 – 30 % of the demand
in industrialized countries. Even the use of purification processes (see page 24) does not reduce the NaCl content below 0.03 wt %. In addition to the NaCl and NaClO3 , the levels of Si,
Ca, Mg, and sulfate impurities are higher than
for the mercury and membrane processes.
9.2. Economics
The wide variation in the main cost factor, that
for electrical energy, which varies from region
to region by a factor of up to three, makes a
direct comparison of production costs problematic. Further, the cost of electrical energy is increasing in different regions at drastically different rates, depending on the basic source of
energy and customs. Rapidly changing foreign
exchange rates also make international comparisons difficult.
A detailed discussion of the capital investment and operating costs for the three processes
for a 200 000 t/a-plant in 1991 is given in [185].
A comparison of the investment costs does not
make sense today for the mercury process, because no mercury cell plant and only a few diaphragm cell plants were built since then. All
new plants are using the membrane process.
9.2.1. Equipment
The expenses for the rectifier, chlorine and hydrogen systems, HCl system, caustic storage,
utilities, and engineering and construction overheads are approximately the same for the three
processes.
Chlorine
65
Table 20. Product qualities: typical compositions of chlorine, caustic, and hydrogen
Product and contents
Process
Diaphragm
Unpurified
Chlorine gas (from cells),
vol%
Cl2
O2
CO2
H2
N2
NaOH solution, wt %
NaOH
NaCl
Na2 CO3
Na2 SO4
NaClO3
SiO2
CaO
MgO
Al2 O3
Fe
Ni
Cu
Mn
Hg
NH3
Hydrogen gas, vol%
H2
Mercury
Membrane
98 – 99
0.1 – 0.3
0.2 – 0.5
0.1 – 0.5
0.2 – 0.5
97 – 99.5
0.5 – 2.0
50.0
0.005
0.05
0.0005
0.0005
<0.001
0.001
0.0002
0.0005
0.0005
50.0
0.005
0.04
0.0001
0.001
0.002
0.0001
0.0001
0.0001
0.0004
Purified
96.5 – 98
0.5 – 2.0
0.1 – 0.3
0.1 – 0.5
1.0 – 3.0
0.03 – 0.3
50.0
1.0
0.1
0.01
0.1
0.02
0.001
0.0015
0.0005
0.0007
50.0
0.025
0.1
0.01
0.001
0.02
0.001
0.0015
0.0005
0.0007
0.0002
0.0002
0.00001
0.0001
none*
none*
0.001
0.00001
none*
>99.9
>99.9**
>99.9
−6
* < 10 %.
** Hydrogen gas from the mercury process contains mercury: 1 µg/m3 – 10 mg/m3 , depending on the purification process. The hydrogen
gas from the other two processes is free of mercury.
Cells. The complex mercury cells are considerably more expensive than the simpler diaphragm and membrane cells. There is no development in mercury cell technology. Improvements are being made in diaphragm cells (higher
current densities, longer service times), but the
relative advantage of the membrane cells is rising fast with considerable increase in current
density and improved membrane performance:
Fewer cells are needed for a given production
capacity.
Brine System. The brine system for the diaphragm process is the simplest of the three –
there is neither sulfate precipitation nor dechlorination – and makes up only 3 – 4 % of the capital
investment. The brine system is the most complex for the membrane process, for fine purification by ion exchange is necessary. However,
the two- or three-fold greater depletion of the
brine in the membrane process allows the brine
system to be smaller than that for the mercury
process. Therefore, the cost of the brine system
for either process is approximately the same, 4
– 7 % of the total.
Caustic Concentration. The elaborate multistage evaporators required for the concentration
of the diaphragm-cell caustic and the separation of NaCl and Na2 SO4 must be nickel plated
because of the corrosiveness of the cell liquor
containing NaCl and NaClO3 . These evaporators cost 20 – 35 % of the total. The evaporators
for the membrane process may be constructed of
stainless steel and are much smaller because the
essentially salt-free cell liquor is more concentrated, costing 3 – 4 % of the total. The mercury
process produces 50 % caustic directly, evaporation is not required.
Facilities for Handling Salt. The mercury
and membrane plants require storage and handling facilities for solid salt. If a diaphragm plant
uses well brine, only small facility is needed for
the recycling of the salt from the caustic evaporation.
66
Chlorine
Mercury. In addition to the capital cost of
mercury itself, there is the expense of the equipment to prevent emission of mercury into the
environment and to remove mercury from the
products (see page 25). This equipment costs
10 – 15 % of the total capital investment.
The investment cost of a new (green-fields)
chlor-alkali project in the USA is estimated to
be between 250 000 and $ 300 000 per tonne per
day chlorine capacity in 1998 [186].
9.2.2. Operating Costs
The fixed costs for operators and other personnel, taxes, insurance, repairs, and maintenance
are about the same for all three processes. The 20
% lower depreciation of the membrane process
is offset by the additional expense for purchase
and replacement of the membranes and for the
more elaborate brine purification.
Of the variable costs, the expense for salt,
precipitants, and anode reactivation are roughly
the same. The difference among the three processes shows up in the consumption of energy,
as electricity and steam. If 1 t of steam is taken
to be equivalent to 400 kWh of electrical energy,
then the comparison in Table 21 can be made.
The differing total energy consumptions are illustrated in Figure 62.
Figure 62. Relative consumption of energy (electricity
and steam) in the three chlor-alkali processes in producing
50 wt % NaOH
The price of electrical energy varies widely
from region to region. The relatively broad range
of possible current densities combined with the
steep increase in the cell voltage with current
density for the diaphragm and membrane cells
allows optimization of the current density with
respect to the local energy price. That is, if electrical energy is relatively expensive, a greater
number of cells, and thus a greater capital investment, can be tolerated to reduce the specific
energy consumption and thus minimize total unit
production cost [187].
9.2.3. Summary
In spite of the advantages of the membrane technology, about 75 % of all chlorine is produced
in mercury and diaphragm cells, operating in ca.
500 plants around the world. Diaphragm technology prevails in the United States (70 %), Russia, and China, and mercury technology in Western Europe (64 %). Continued production from
these plants is economical under special circumstances.
For mercury cell users, the question of today
is whether the old, depreciated plant is competitive with new membrane cell plants. The alternatives are:
– Further production in the mercury cell plant
– Conversion to the membrane process
– Phasing out the old plant
Candidates for further production are plants
of medium to large size, with low electricity
costs, with very high quality products, with
high emissions standards, with high maintenance standards (low repair costs), or which
produce speciality products which cannot be
obtained in membrane cells (e.g., alkoxides or
dithionites).
All producers that do not fulfill one or more
of these conditions are candidates for conversion. The more the existing infrastructure can
be used, the greater the benefits resulting from
conversion. The investment for the conversion
of a middle-sized plant (100 000 t/a) is between
$ 550 and $ 800 per tonne of chlorine capacity
per year [188]. This investment includes the cost
for the membrane cells, secondary brine purification and additional changes of the infrastructure. The costs are specific for each existing plant
and depend on:
Chlorine
67
Table 21. Energy consumed to produce 1 t of chlorine plus 1.13 t of caustic soda (50 %) in the three chlor-alkali processes
Energy
Electricity for
electrolysis, kWh
Steam equivalent, kWh
Process
Diaphragm
Mercury
Membrane
2300 – 2900
3100 –
3400
2200 –
2600
800 – 1000
Total, kWh
3100 – 3900
Relative
energy costs
100 %
0
3100 –
3400
93 %
200 – 400
2400 –
3000
78 %
– Chlorine quality (e.g., the oxygen content)
– The use of existing buildings. The materials
of the existing brine treatment area
– The possibility of using the dilute caustic
(32 – 36 %) within the plant without concentration
– Use of the existing electrical equipment, rectifiers, busbars
– Possible capacity enlargement because of the
lower specific energy per tonne of chlorine
– Dismantling and disposal of mercurycontaminated parts of the old plant.
a plant with a capacity of 1000 t/d to membrane
technology is ca. 90 million dollars [186].
In the first few years after the introduction
of membrane technology, diaphragm cells in
several plants were equipped with membranes
(retrofit) to reduce the cost of steam for cell
liquor concentration, to give a small reduction
in electricity consumption and better quality
of caustic. This procedure is economic where
steam is very expensive [187, 191].
Normally the decision for a conversion is initiated by plans for an expansion of the production capacity or by environmental legislation.
Each change in the plant structure or in the cost
structure may lead to reevaluation of the future of
the electrolysis plant. Therefore, each plant has
to be considered individually [189, 190]. For the
European chlor-alkali industry a detailed analysis of the impact of a conversion of all mercury
cells to the membrane technology on the competitiveness of the industry is given in [110].
The situation is different for diaphragm cell
plants. These plants are still economic where
inexpensive brine (e.g., from solution mining)
is available, energy costs are comparably low
(e.g., from cogeneration of electricity and steam
on site), and when the market price for caustic is determined by the lower quality of diaphragm caustic. In countries like the United
States it will be difficult to economically justify
conversion. In contrast to the mercury process,
improvements to the cells are still being made,
resulting in lower operating costs and savings in
solid waste disposal. The investment to convert
10. Other Production Processes
10.1. Electrolysis of Hydrochloric Acid
Electrolytic decomposition of aqueous hydrochloric acid is used to produce chlorine and
hydrogen. The first pilot plant was set up by
G. Messner in 1942 in Bitterfeld, Germany, and
since 1964 eight full-scale plants have been commissioned in Europe and the United States, a
total capacity of 540 000 t/a [192]. Hydrogen
chloride is a byproduct of many organic industrial processes. Electrolysis of hydrochloric acid
competes with chemical processes in which either hydrogen chloride is used to produce chlorinated hydrocarbons directly, e.g., by oxychlorination, or where chlorine is produced by chemical reaction, e.g., in the KEL chlorine process
(see page 69). The advantages of the electrolytic
process are very pure products without further
treatment, reliability (simple design), ease of operation, flexibility (5 : 1 turndown ratio), and low
energy consumption even with small installations.
68
Chlorine
Principles. Hydrochloric acid (22 wt % HCl)
is fed into the cells in two separate circuits, a
catholyte circuit and an anolyte circuit. During
electrolysis the concentration is reduced to ca.
17 %, and the temperature increases from 65 to
80 ◦ C. A part of the depleted acid is separated
from the catholyte stream, concentrated in the
absorption plant to ca. 30 %, and fed back into
the main stream. The electrolyzer is bipolar, with
pairs of electrodes arranged like the leaves of
a filter press. A diaphragm or membrane (e.g.,
Nafion 430) separates the anode compartment
from the cathode compartment to prevent mixing of the gaseous products.
The reversible standard decomposition potential of hydrochloric acid is 1.358 V, made up
of the anode potential, the discharge of chloride
ions with formation of chlorine, and the cathode
potential, the discharge of hydroxonium (H3 O+ )
ions with formation of hydrogen. In practice
(> 15 % HCl, 70 ◦ C), the decomposition potential is ≤ 1.16 V.
The graphite electrode plates are not attacked
by 22 % hydrochloric acid. A poly(vinyl chloride) (PVC) fabric constitutes the diaphragm.
Chlorine dissolved in the anolyte diffuses
through the diaphragm and is reduced at the
cathode, causing a loss of 2 – 2.5 % of the theoretical current yield. The increase of cell voltage when current flows is mainly because of
the hydrogen overpotential at the graphite cathode and the resistance of the electrolyte. Depolarizing agents (polyvalent metal ions) in the
catholyte reduce the overpotential by ≤ 300 mV
at 4 kA/m2 [193].
The conductivity of hydrochloric acid is maximized at a concentration of 18.5 wt %. High
temperatures improve the conductivity, but to
avoid increased vapor pressure of HCl and material problems, the temperature is kept below
85 ◦ C. Modern cells have a voltage of ca. 1.90 V
at 4.8 kA/m2 , corresponding to an energy consumption of 1400 – 1500 kWh per tonne of chlorine.
Diaphragm Cells. Hydrochloric acid electrolysis cells are manufactured by Hoechst –
Uhde [194]. Each Hoechst – Uhde electrolyzer
consists of 30 – 36 individual cells that are
formed from vertical graphite plates connected
in series, between which there are diaphragms.
To improve gas release, vertical slits are milled
in the graphite plates, which are cemented in
frames made of HCl-resistant plastics. At the
bottom of the frames, channels feed in the electrolyte. The gases rise up the plates and pass
through ducts into collection channels in the
upper part of the cell. Chlorine leaves the cell
with the anolyte, and hydrogen leaves with the
catholyte. The end plates of the electrolyzer are
made of steel lined internally with rubber and
are held together by spring-loaded tension rods.
The electric current is supplied via graphite terminals. The unit rests on insulated steel frames.
The effective surface of the electrodes is 2.5 m2 ,
and the current loading can be up to 12 kA.
DeNora and General Electric are developing
an electrolyzer with a solid polymer electrolyte
(SPE) based on Nafion [195]. In addition to a
voltage savings of 20 %, it is hoped that completely chloride-free hydrogen gas can be produced.
Operation. A simplified flow diagram of the
process as operated by Bayer – Hoechst – Uhde
is shown in Figure 63.
Figure 63. Simplified flow diagram of a hydrochloric acid
electrolysis
a) Absorption column; b) Heat exchanger; c) Strong acid
tank; d) Catholyte collecting tank; e) Catholyte filter;
f) Catholyte supply tank; g) Electrolyzer; h) Hydrogen –
catholyte separator; i) Chlorine – anolyte separator;
k) Anolyte collecting tank; l) Anolyte filter; m) Anolyte
supply tank; n) Weak acid line to absorber
In the absorption column, the hydrogen chloride gas is absorbed adiabatically by depleted
hydrochloric acid from the catholyte. In the upper section of the column, an absorber removes
the remaining hydrogen chloride and the water
vapor by absorption in a water stream, which
Chlorine
makes up the water balance of the process. The
30 wt % acid that is produced is then cooled,
purified if necessary by activated carbon, and
supplied to the anolyte and catholyte circulation
systems.
The electrolyte is pumped through a filter and
heat exchanger to a gravity feed tank for the
electrolyzer unit. The gases produced are freed
from the electrolytes in separators, and the electrolytes flow back into their respective collecting
tanks to be resaturated. The working life of the
PVC diaphragms, 1 – 2 years, depends on the
impurities in the acid. The concentrated acid is,
therefore, purified carefully [196].
The product gases are saturated with water
vapor and hydrogen chloride at the partial vapor pressures of 20 % hydrochloric acid. Both
product streams are cooled. Sodium hydroxide
solution is used to wash the hydrogen, removing
chlorine and hydrogen chloride and producing a
99.9 % product. The chlorine, which is dried by
sulfuric acid, contains ca. 0.5 % hydrogen and
ca. 0.05 % carbon dioxide. The hydrogen overpotential can be reduced by activation of the
cathodes.
Membrane Cells [197, 198]. Since 1992
Bayer has replaced the woven fabric cloth in the
diaphragm cells by anion-exchange membranes
of the sulfonate type. Only hydrated protons are
able to pass from the anolyte to the catholyte, so
that the whole cell and the electrolyte systems
are simplified. Together with an optimized surface of the electrodes for better gas release, this
leads to:
– Lower cell voltage of 300 mV, corresponding to a power consumption of 1300 kWh per
tonne of chlorine at 4.8 kA/m2
– Longer life of the cell components
– Higher product quality
– Improved safety of operation
– Simplified process
A similar electrolytic process for recovering
chlorine from anhydrous HCl, also using membrane cell technology, has been developed by
DuPont [199].
10.2. Chemical Processes
The chlor-alkali process produces chlorine and
sodium hydroxide solution in fixed stoichiomet-
69
ric proportions. Experience has shown that there
tends to be a surplus of either chlorine or sodium
hydroxide. Chlorine may, however, be produced
competitively without the byproduct sodium hydroxide by nonelectrolytic methods. The starting
material is usually hydrogen chloride, which is
catalytically oxidized to chlorine by oxygen, air,
nitric acid, sulfur trioxide, or hydrogen peroxide.
Other processes start from ammonium chloride
or metal chlorides.
10.2.1. Catalytic Oxidation of Hydrogen
Chloride by Oxygen
A catalyst is essential for the economic oxidation of hydrogen chloride to chlorine by air or
oxygen (Deacon Process), and the catalyst must
be active at low temperature and have adequate
life. There are many patents claiming improved
catalysts and equipment. Most of the catalysts
are oxides and/or chlorides of metals on various substrates. Only three processes have been
commercialized.
The KEL Chlorine Process. The process
developed by Kellogg [197] uses concentrated
sulfuric acid (ca. 80 %) with ca. 1 % nitrosylsulfuric acid as the catalyst. From 1975 to 1988
Du Pont operated a full-scale plant in Corpus
Christi, Texas, recovering up to 600 t/d of chlorine. The plant was shut down due to a change
in the structure of the plant and because of
material problems after more than 10 years of
operation. The raw material, from a fluorinated
hydrocarbon plant, consisted of waste gases that
contained hydrogen chloride [200]. Figure 64
shows a simplified flow diagram.
Sulfuric acid catalyst is fed into the top of
the stripper column. The hydrogen chloride gas
reacts with the catalyst to form nitrosyl chloride:
HCl+NOHSO4 → NOCl+H2 SO4
The oxygen, the ultimate oxidizing agent,
blows the remaining hydrogen chloride out of
the sulfuric acid, which becomes more concentrated and also is cooled in a flash vaporizer. This
acid is then fed back into the process. Nitrosyl
chloride, hydrogen chloride, oxygen, and water
70
Chlorine
vapor flow as a gaseous stream into the oxidizer
and react there, increasing the temperature:
2 NOCl+O2 → 2 NO2 +Cl2
The cooled, dried chlorine gas still contains
ca. 2 % hydrogen chloride and up to 10 % oxygen. Both are removed by liquefaction.
The net reaction is
NO2 +2 HCl → NO+Cl2 +H2 O
4 HCl+O2 → 2 Cl2 +2 H2 O
In the absorber – oxidizer, the rest of the hydrogen chloride is oxidized. Concentrated sulfuric acid is fed in at the top, reacts with the oxides
of nitrogen to form nitrosylsulfuric acid, absorbs
the water that has formed, and is conducted back
into the stripper:
NO+NO2 +2 H2 SO4 → 2 NOHSO4 +H2 O
NOCl+H2 SO4 → NOHSO4 +HCl
Figure 64. Flow diagram of the KEL chlorine process (simplified)
a) Stripper; b) Oxidizer; c) Absorber – oxidizer; d) Acid
chiller; e) Acid cooler; f) Vacuum flash evaporator
The installation at Corpus Christi operated at
1.4 MPa and 120 – 180 ◦ C. On account of the aggressive nature of the chemicals, expensive materials, such as tantalum-plated equipment and
pipes, must be used. For outputs of 250 – 300 t
of chlorine per day, this process can be more
economical than the electrolysis of hydrochloric acid, depending on local conditions.
The Shell Chlorine Process. The catalyst
developed by Shell consists of a mixture of copper(II) chloride and other metallic chlorides on
a silicate carrier [201]. The reaction of the stoichiometric mixture of hydrogen chloride and
air takes place in a fluidized-bed reactor at ca.
365 ◦ C and 0.1 – 0.2 MPa. The yield is 75 %. The
water condenses out from the gas stream, and
the hydrogen chloride is removed by washing
with dilute hydrochloric acid. After the residual
gas has been dried with concentrated sulfuric
acid, the chlorine is selectively absorbed, e.g.,
by disulfur dichloride. After desorption and liquefaction, the chlorine has a purity > 99.95 %.
A manufacturing unit was built by Shell in
the Netherlands, 41 000 t/a, and another in India, 27 000 t/a, but both have been closed down
owing to the prolonged surplus of chlorine on
the market.
The Mitsui MT-Chlorine Process. The catalyst consists of chromium(III) oxide on a silicate carrier [202]. In a fluidized-bed reactor,
hydrogen chloride is reacted with oxygen at a
temperature of 415 ◦ C to give chlorine gas with
a conversion rate of 73 – 77 %. The reactor is
made from nickel-lined low carbon steel. The
concentration of the purified product is > 99.5
% Cl2 .
A commercial plant for 30 000 t Cl2 /a is successfully operating since 1988, with an expansion to 60 000 t Cl2 /a in 1990.
Chlorine
A two-stage cyclic fluidized bed process for
converting HCl to chlorine is described in [203].
The catalytic oxidation process combines the
exothermic oxidation of 60 – 70 % of the HCl at
380 – 400 ◦ C in a fluidized bed of copper oxychlorides impregnated on zeolite with the transfer of the reaction products to a second reactor
operating at 180 – 200 ◦ C where the rest of HCl
is converted.
71
vapor at high temperature and may also contain brine mist and traces of chlorinated hydrocarbons, and is normally at atmospheric pressure. Before the chlorine can be used, it must be
cooled, dried, purified, compressed, and where
necessary, liquefied. A simplified flow sheet is
shown in Figure 65.
11.1. Cooling
10.2.2. Oxidation of Hydrogen Chloride by
Nitric Acid
The nitrosyl chloride route to chlorine is based
on the strongly oxidizing properties of nitric
acid:
6 HCl+2 HNO3 → 2 Cl2 +2 NOCl+4 H2 O
2 NOCl+2 H2 O+O2 → 2 HCl+2 HNO3
The practical problems lie in the separation of
the chlorine from the hydrogen chloride and nitrous gases. The dilute nitric acid must be reconcentrated. Corrosion problems are severe. Suggested improvements include (1) oxidation of
concentrated solutions of chlorides, e.g., LiCl,
by nitrates followed by separation of chlorine
from nitrosyl chloride by distillation at 135 ◦ C
or (2) oxidation by a mixture of nitric and sulfuric acids with separation of the product chlorine
and nitrogen dioxide by liquefaction and fractional distillation [204].
10.2.3. Production of Chlorine from
Chlorides
Alkali-metal chlorides, ammonium chloride,
and other metallic chlorides are reacted, usually with nitric acid, to produce nitrate fertilizers
[205]. Chlorine is not produced directly, but it
can be obtained from the intermediate products
nitrosyl chloride or hydrochloric acid.
11. Chlorine Purification and
Liquefaction
Chlorine produced by the various processes, especially by electrolysis, is saturated with water
Table 22 shows the volume, water content, and
heat content of 1 kg of chlorine gas at 101.3 kPa
as a function of temperature. To avoid solid chlorine hydrate formation, the gas is not cooled below 10 ◦ C [206]. Cooling is accomplished in either one stage with chilled water or in two stages
with chilled water only in the second stage.
Table 22. Volume, moisture content, and enthalpy of 1 kg of
chlorine gas at 101.3 kPa as a function of temperature t
t, ◦ C
Volume, m3
Dry
Saturated *
20
40
0.312
0.335
0.357
0.314
0.342
0.385
60
70
80
0.380
0.392
0.404
0.473
0.565
0.756
0
Water
Heat
content,
g/kg **
content,
kJ/kg *
1.54
5.95
3.81
24.45
69.50
19.7
61.5
112.0
222.0
188.41
325.73
623.83
* Chlorine gas saturated with water vapor at temperature t.
** Grams of H2 O per kg of Cl2 .
The chlorine gas can be cooled indirectly in a
tubular titanium heat exchanger so that the cooling water is not contaminated and the pressure
drop is small. The resultant condensate is either
fed back into the brine system of the mercury
process or dechlorinated by evaporation in the
case of the diaphragm process.
The chlorine gas can be cooled directly in
packed towers. Water is sprayed into the top
and flows countercurrent to the chlorine. This
treatment thoroughly washes the chlorine; however, dechlorination of the wastewater consumes
a large amount of energy. The cooling water
should be free of traces of ammonium salts to
avoid the formation of nitrogen trichloride.
Closed-circuit direct cooling of chlorine
combines the advantages of the two methods.
72
Chlorine
Figure 65. Simplified flow diagram of a chlorine processing plant
a) Chlorine gas cooler (primary); b) Chlorine demister; c) Blower or fan; d) Chlorine gas cooler/chiller (secondary);
e) Condensate collection tank; f) Drier, first stage; g) Drier, second stage; h) Sulfuric acid mist separator; i) Sulfuric acid
circulation pump; k) Cooler for circulating sulfuric acid; l) Sulfuric acid feed tank; m) Cooler for sulfuric acid feed
The chlorine-laden water from the cooling tower
is cooled in titanium plate coolers and recycled.
The surplus condensate is treated like the condensate from indirect cooling. Spray towers, as
well as packed towers, are used. Water carryover is removed by demisters, which reduce the
amount of sulfuric acid used for drying.
11.2. Chlorine Purification
Water droplets and impurities such as brine mist
are mechanically removed by special filter elements with glass wool fillings. The efficiency
varies with the gas throughput. A commonly
used device is the Brink demister [207]. Instead
of glass wool, porous quartz granules can be
used.
In electrostatic purification, the wet chlorine
gas is passed between wire electrodes in vertical tubes. The electrodes are maintained at a
d.c. potential of 50 kV with a current density
of 0.2 mA/m2 . The particles and droplets in the
chlorine become charged and collect on the tube
walls. The resultant liquid is fed back into the
brine system or chemically treated before disposal.
Activated carbon filters can adsorb organic
impurities and may be regenerated by heating to
200 ◦ C.
Gaseous impurities can be removed by absorption of the chlorine in a suitable solvent,
such as carbon tetrachloride, water, or disulfur
dichloride, followed by desorption. This can be
coupled with further processes, such as the recovery of chlorine from the waste gas remaining
after liquefaction [15, pp. 418 – 422].
A wash with concentrated hydrochloric acid
removes the dangerously explosive nitrogen
trichloride [208]. Scrubbing with liquid chlorine
(see Fig. 67) mainly reduces the content of organic impurities and carbon dioxide, but it can
also lower the bromine content. When the chlorine is cooled down to near its dew point, liquid
chlorine scrubbing is often combined with compression by turbo or reciprocating compressors.
11.3. Drying
Drying of chlorine is carried out almost exclusively with concentrated sulfuric acid (96 –
98 wt %) [209]. Depending on the desired final concentration of the waste acid, drying can
be a two-, three-, or four-stage process. The
acid and chlorine flow countercurrently. The fi-
Chlorine
nal moisture content depends on the concentration and temperature of the acid in the final stage (Fig. 66). An upper limit is 50 ppm
H2 O. Low-temperature liquefaction (− 70 ◦ C)
demands lower moisture content, which can be
achieved with molecular sieves, whereby 2.5
ppm is possible [210].
73
sion from increasing the temperature enough to
ignite material in contact with the chlorine.
Wet chlorine gas can be compressed 20 –
50 kPa by a single-stage blower or fan with a
rubber-lined steel casing and titanium impeller.
It can also be compressed in liquid-ring compressors, so that further treatment of the chlorine can be accomplished in smaller equipment
[211]. Sulfuric acid ring compressors are used
for throughputs of 150 t of dry chlorine gas per
day per compressor and for pressures of 0.4 MPa
or, in two-stage compressors, 1.2 MPa. The heat
of compression is removed by cooling the circulating liquid; cooling of the gas is not necessary. Advantages are simplicity of construction, strength, and reliability, but efficiency is
low [212].
Reciprocating compressors were formerly lubricated with sulfuric acid, but are now available
as dry-ring compressors (no lubrication). They
can compress up to 200 t per day. Multistage
compressors produce pressures up to 1.6 MPa.
The heat of compression of each stage must be
removed by heat exchangers or by injection of
liquid chlorine (see Fig. 67). Well-purified chlorine gas is essential for trouble-free operation
[213].
Figure 66. Drying chlorine with sulfuric acid
Attainable moisture content as a function of concentration
and temperature of the acid
The packed towers usual in the first stages
are constructed of rubber-lined steel or glassfiber-reinforced poly(vinyl chloride). The heat
liberated on dilution of the circulating acid is
removed by titanium heat exchangers, and the
weak acid is dechlorinated chemically or by
blowing air. Often the acid is recirculated after
reconcentration to 96 % by heating under vacuum. Generally, columns with bubble cap plates
or sieve trays are used at the final stage. The drying is effective, but the pressure drop is great.
Occasionally, spray towers are used to dry chlorine.
After drying, the chlorine gas is passed
through a demister or a packed bed to remove
sulfuric acid mist.
11.4. Transfer and Compression
In all operations involving compression, care
must be exercised to prevent the heat of compres-
Figure 67. Multistage reciprocating compressor for chlorine liquefaction at 1 MPa with cooling water at 15 ◦ C with
liquid chlorine scrubbing
a) Low-temperature cooling and scrubbing column;
b) Collection tank for impurities; c) Three-stage compressor; d) Intermediate cooler, stage 1; e) Intermediate
cooler, stage 2; f) Liquefier; g) Chlorine collection vessel;
h) Chlorine storage tank; i) Chlorine storage tank on load
cells
Turbo compressors are most economical
when they operate with large amounts of chlorine. Each unit compresses up to 1800 t/d.
In multiple-stage operation, pressures up to
74
Chlorine
1.6 MPa are reached. Labyrinth seals are used
on the high-speed shafts. Requirements for cooling and gas purity are like those of reciprocating compressors. Screw compressors handle
low rates of chlorine and give pressures up to
0.6 MPa. Sundyne blowers are one-stage highspeed centrifugal compressors handling 80 –
250 t per day and giving pressures up to 0.3
MPa. Liquid chlorine injection is used for cooling [214]. Membrane compressors are used for
pressurizing storage tanks with chlorine gas to
transfer liquid chlorine to other vessels [215].
Liquid chlorine is pumped with canned pumps
[216].
11.5. Liquefaction
The most suitable liquefaction conditions can
be selected within wide limits. Important factors are the composition of the chlorine gas, the
desired purity of the liquid chlorine, and the desired yield. There are nomograms that give the
relationship between the chlorine concentrations
of the incoming and residual gases, liquefaction
yields, pressures, and temperatures [217]. Increasing the liquefaction pressure increases the
energy cost of chlorine compression, although
the necessary amount of cooling decreases, resulting in an overall reduction in energy requirement (Table 23) [218].
Table 23. Electrical energy requirement for compression and
liquefaction of 1 t of chlorine gas
Liquefaction pressure, MPa
0.1
0.3
0.8
1.6
Energy for compression,
kWh/t
Energy for cooling, kWh/t
Combined energy, kWh/t
5
87
92
23
68
91
42
27
69
57
3
60
Starting temperature, ◦ C
Final temperature, ◦ C
−36
−42
−8
−17
25
14
53
40
Any hydrogen is concentrated in the residual gas. To keep the hydrogen concentration below the 6 % explosive limit, conversion of gas to
liquid should be limited to 90 – 95 % in a singlestage installation. Higher yields may be obtained
by condensing the chlorine from the residual
gas in a second stage, which is constructed to
reduce the risk from explosion [219]. This is
achieved by the use of sufficiently strong equipment to withstand explosions or by the addition
of enough inert gas to keep the mixture below
the explosive limit. Multistage installations can
liquefy over 99.8 % of the chlorine gas.
High-pressure (0.7 – 1.6 MPa) liquefaction
with water cooling (Fig. 67) does not require a
cooling plant. Therefore, it has the lowest energy
cost of all methods; however, the high construction cost must be set against this.
Medium-pressure (0.2 – 0.6 MPa) liquefaction with cooling (− 10 to − 20 ◦ C) is especially
useful when only a part of the chlorine is to be
liquefied and the remaining gas is to be reacted
at the liquefaction pressure, e.g., with ethylene
to form ethylene dichloride. The residual gas can
be fed into the compressor suction systems, provided that the increased inert gas content does
not interfere with the subsequent process. Otherwise, the residual gas must be scrubbed free
of chlorine or liquefied in a second stage.
Figure 68 shows a two-stage liquefaction
by the Uhde system, which operates at 0.3 –
0.4 MPa and − 20 ◦ C in the first stage and
− 60 ◦ C in the second stage, with a yield
of 99 % [220]. The refrigerant is difluoromonochloromethane. The gaseous refrigerant
is compressed, liquefied by water cooling, and
collected in a container. The liquid refrigerant is
sprayed into the shell of the chlorine liquefier,
where it evaporates, absorbing heat and cooling the chlorine, which flows from the liquefiers
at − 15 ◦ C (first liquefier) or − 55 ◦ C (second
liquefier) [221]. The residual gas from the first
horizontal liquefier contains < 5 % hydrogen. It
is fed into the second liquefier, which is at an
angle of 60◦ and has a strong, low-volume construction. There the gas mixture passes through
the explosive concentration limits. In case of an
explosion, there is a comprehensive control system to ensure safety:
The explosion pressure is vented by means of
a bursting disk to a residual gas absorber. Simultaneously the residual gas from the first stage is
passed directly into this absorber. The chlorine
gas to the second stage is shut off, and an inert
gas purge is introduced. Finally, the liquid chlorine exit valve is closed to prevent back flow of
the liquid chlorine into the second liquefier and
from there into the absorber.
With normal-pressure (ca. 0.1 MPa) liquefaction and low temperature (< − 40 ◦ C), cryogenic storage of the liquid chlorine is possible.
Chlorine
75
Figure 68. Flow diagram of a two-stage chlorine liquefaction plant at intermediate pressure – Uhde system
a) Chlorine gas compressor; b) Refrigerant collector, stage 1; c) Refrigerant condenser, stage 1; d) Chlorine liquefier, stage
1; e) Refrigerant separator, stage 1; f) Refrigerant compressor, stage 1; g) Liquid chlorine storage tank; h) Chlorine liquefier,
stage 2; i) Bursting disk; j) Refrigerant separator, stage 2; k) Refrigerant condenser, stage 2; l) Refrigerant collector, stage 2;
m) Refrigerant compressor, stage 2
This process is advantageous when large quantities of chlorine must be liquefied as completely
as possible. Attention must be paid to the increased solubility of other gases at low temperatures, especially carbon dioxide [206]. This
carbon dioxide can be removed from the liquid
chlorine by passage of hot chlorine gas [219].
An absorption – desorption process by Akzo
is based on carbon tetrachloride [222]. It requires
little energy and yields over 99.8 % of a pure liquid chlorine that is almost free of carbon dioxide. A similar process by Diamond Shamrock
has been described [224].
11.6. Chlorine Recovery
Chlorine can be recovered from the tail gas from
liquefaction with a chlorine recovery system.
Tail gas from liquefaction and chlorine from
the plant evacuation system together with the
snift compressor and stripper recycle streams are
supplied to a snift compressor suction knockout drum. The gas is compressed by the snift
gas compressor to 7.0 kg/cm2 with a discharge
temperature of 85 ◦ C.
The snift gas is then cooled by cooling water
to 45 ◦ C and then further cooled to − 12.2 ◦ C
by Freon. Gas is sent to the absorber, whereas
liquid is either returned to chlorine storage or is
used for reflux at the stripper.
The off-gas enters the bottom of the chlorine absorber and passes upward through the two
packed sections of the tower while cold carbon
tetrachloride flows downward. All of the chlorine and the nitrogen trichloride is absorbed in
the carbon tetrachloride while the noncondensable gases remain in the gasphase and are removed from the system.
The chlorine-rich carbon tetrachloride leaves
the bottom of the chlorine absorber at ca. 10 ◦ C
and is forced by pressure difference to the chlorine stripper. Chlorine stripper feed enters the
middle of the column and flows downward
through two packed sections, releasing chlorine
as it is heated. A thermosiphon reboiler is provided at the base of the stripper. By heating
the liquid above 65 ◦ C, the absorbed nitrogen
trichloride decomposes to nitrogen and chlorine
gas.
Chlorine boiled off in the stripper passes upward through a packed top section of the column where it is scrubbed and purified by liq-
76
Chlorine
uid chlorine from the discharge knock-out drum.
The stripper overhead stream, a mixture of chlorine and a small amount of inerts, is sent to the
chlorine liquefaction system or recycled to the
suction knock-out drum to maintain the stripper
reflux [222, 223].
12. Chlorine Handling
Both the chlorine industry and governmental organizations are well aware of the risks of chlorine. In the United States and Canada, The Chlorine Institute [225] has established standards and
recommendations for safe transport and handling of chlorine since 1924. In Europe, Euro
Chlor, an association of major Western European chlorine manufacturers, publishes recommendations, codes, and memorandums for chlorine handling and transport concerning European conditions and regulations [226]. Both organizations distribute manuals and pamphlets
worldwide. Surveys of existing national and
international regulations for the handling and
transport of hazardous chemicals are available
[2, 34, 227, 229].
shell around a double-enveloped low-pressure
storage tank can provide such a facility. To vent
chlorine, there must be an absorption or liquefaction system. In the course of all operations,
the design pressure should not be exceeded. The
dimensions of branches and the amount of pipe
work should be minimized. Bottom connections
from storage tanks are not recommended for
small chlorine users. Large branches should always be located in the gas space of a vessel.
The pipework system should be provided with
remotely operable valves to permit isolation in
case of emergency.
Before being put into service, the whole storage system must be degreased, cleaned, and
dried to achieve a dew point of − 40 ◦ C in the
purge gas at the outlet of the system. No substance that could react with the chlorine can be
allowed to enter the storage system. The filling
ratio in the tank should never exceed 95 % of the
total volume of the vessel; for pressure storage
tanks, this corresponds to 1.25 kg of liquid chlorine per liter of vessel capacity at 50 ◦ C (Fig.
69).
12.1. Storage Systems
Chlorine is liquified and stored at ambient or
low temperature [230, 233]. In both cases the
pressure in the storage system corresponds to
the vapor pressure of liquefied chlorine at the
temperature in the stock tank. Pressure storage
is recommended for all usual customers [234,
235]. Euro Chlor recommends a maximum capacity of 300 – 400 t for individual tanks. For the
large storage capacities required by producers,
usually a low-pressure storage system, operating at a liquid chlorine temperature of ca. − 34
◦
C, is chosen. A low-pressure system needs a
cooling or recompression system, and, for this
reason, it is basically unsuitable for small chlorine consumers [236, 237].
A few major design aspects must be mentioned. Any risk of fire or explosion must be
eliminated. All tanks having an external connection below the liquid level should be placed in a
liquor-tight embankment (bund). In the event of
leakage the liquid should be collected in a small
area to reduce the rate of vaporization. The outer
Figure 69. Proportion (%) of a one-liter vessel occupied
by 1.25 kg of liquid chlorine as a function of temperature
Typical measuring and control equipment of
a pressure storage tank is shown in Figure 70.
The ISO codes for process measurement control
functions and instrumentation are explained in
Table 24 [238]. The measuring equipment of a
low-pressure storage system needs supplementary devices, for example, a temperature indicator with an alarm and, in the case of a double-
Chlorine
shell vessel, a device to determine the quality
of the purging gas inside the double shell. The
vessel and an external envelope should be protected against overpressure or underpressure. In
low-pressure systems, the chlorine is removed
by vertical submerged pumps, canned pumps below the vessel, or ejector pumps operating with
a flow of liquid chlorine produced by external
pumps.
Table 24. ISO codes and miscellaneous symbols for process
measurement control functions and instrumentation
Codes
Function or Instrumentation
AA
CW
dPI
FA
FI
FIA
FICA
FRA
HZ
H
L
LA
LIA
LIC
M
PA
PCZA
PI
PIA
PIAS
PIC
PRC
PSA
QRA
TA
TC
TI
TIA
TIC
TRA
WI
WIA
analysis alarm
cooling water
difference pressure indicating
flowrate alarm
flowrate indicating
flowrate indicating alarm
flowrate indicating controlling alarm
flowrate recording alarm
hand operated emergency acting
high
low
level alarm
level indicating alarm
level indicating controlling
moisture analysis
pressure alarm
pressure controlling emergency acting alarm
pressure indicating
pressure indicating alarm
pressure indicating alarm switching
pressure indicating controlling
pressure recording controlling
pressure switching alarm
quality recording alarm
temperature alarm
temperature controlling
temperature indicating
temperature indicating alarm
temperature indicating controlling
temperature recording alarm
weight indicating
weight indicating alarm
measuring device
remote control valve
control line
–––
Periodic inspection and retesting of the whole
system, including a visual examination, a thickness test of the wall of the vessel and pipes, and
an examination of the welds and the surfaces
under any thermal insulation, is recommended.
Hydraulic retesting is accompanied by risk of
corrosion and is, therefore, not favored.
77
12.2. Transport
Within a chemical plant and over distances of
several kilometers, chlorine can be transported
by pipelines, either as gas or liquid [24, 239].
Every precaution should be taken to avoid any
vaporization of chlorine in a liquid-phase system or any condensation in a gas-phase system.
Wherever liquid chlorine could be trapped between two closed valves or wherever the system
could be overpressurized by thermal expansion,
an expansion chamber, a relief valve, or a rupture
disk should be provided [240, 241].
Commercial chlorine is transported as a liquid, either in small containers (cylinders and
drums) or in bulk (road and rail tankers, barges,
and ISO containers). The design, construction,
system of labeling, inspection, and commissioning are covered by national and international regulations [227]. Cylinders have a chlorine content up to 70 kg. A protective hood is provided
to cover the valve during transport. The ton containers (drums) have a capacity of 500 – 3000 kg
of chlorine. Drums are equipped with two valves
near the center of one end and connected with
internal eductor pipes.
The capacity of tank cars (rail tankers)
ranges from 15 to 90 t. Special angle valves are
mounted on the manhole cover on top of the vessel. In Europe, pneumatic valves are normally
used [230 – 232]. During loading and unloading, these valves can be closed rapidly and remotely in case of an accident. They have an internal safety plug, providing a tight seal against
the passage of gas or liquid chlorine in the event
of failure of the body of the valve.
In North America, the eductor pipe inside the
vessel has an excess-flow valve at the top, immediately below the manhole cover. This valve
closes the eductor pipe when the rate of liquid
flow exceeds a set rate [2, 24]. North American tank cars have a spring-loaded safety relief
valve, which protects the vessel against overpressure in case of external heat. The tanks have
thermal insulation. In Europe thermal insulation
and safety relief valves are not used or recommended.
Road tankers and ISO containers have a chlorine capacity of 15 – 20 t. The design of and the
equipment on chlorine pressure road tankers is
similar to these of rail tankers. In North America, large amounts of chlorine are transported by
78
Chlorine
Figure 70. Discharge of liquid chlorine by padding to pressure storage
a) Liquid-chlorine rail tanker; b) Flexible connection; c) Plug; d) Viewing glass; e) Remote-control tank valves; f) Protective
membrane; g) Storage vessel; h) Rupture disk; i) Relief safety valve; j) Buffer vessel for liquid chlorine
tank barges [24]. These barges usually are of the
open-hopper type with several cylindrical uninsulated pressure vessels. The total capacity of
barges ranges from 600 to 1200 t. The chlorine
is transported at low temperature.
Classification and Labeling. According to
Directive EEC 67/548, Annex I, chlorine (Index no. 017–001–00–7) is classified as toxic
and dangerous to the environment. The R and
S phrases are:
R 23
R 36/37/38
R 50
S 7/9
toxic by inhalation
irritating to eyes, respiratory system and skin
very toxic to aquatic organisms
keep container tightly closed and in a
well-ventilated place
For transportation, chlorine is in class 2, no.
2TC (toxic, corrosive) in ADR, RID, and ADNR
in Europe, and in class 2.3 in the IMDG Code (p.
2116) and ICAO Code. All vessels must be labeled with the denomination 268/1017 Chlorine
and with labels for dangerous goods: cylinders
and drums with labels 2.3 (toxic gas) and 8 (corrosive); railroad tankers with labels 2.3, 8, and
13 (shunt carefully); and for shipping, labels 2.3,
8, and MP (marine pollutant).
12.3. Chlorine Discharge Systems
All containers should be discharged in the same
order as received. They must be placed where
any external corrosion, risk of fire, explosion, or
damage is avoided [24]. At normal room temperature, the discharge rate of chlorine gas from
a single 70-kg cylinder is ca. 5 kg/h and the rate
of a drum is ca. 50 kg/h. The flow of chlorine gas
can be increased by a higher ambient temperature or by connecting two or more containers. A
system of two or more containers must be carefully operated and controlled to avoid overfilling
by transfer of chlorine from warm to cool containers. Direct heating of containers is not recommended [24]. The best way to determine the
flow rate and container content is to observe the
weight of the container [228]. A flexible tube is
used to connect a mobile container with the fixed
piping system. Any reverse suction from the consuming plant must be prevented by a barometric leg or other adequate precaution if the chlorination process runs at atmospheric pressure.
Pressurized processes need a pressure controlling system with automatic isolation valves.
Uninsulated tanks have a maximum gaseous
discharge rate of ca. 2 t/h. The chlorine gas can
be used only for low-pressure chlorination processes and at low rates. This method increases
Chlorine
79
Figure 71. Liquid chlorine vaporizer
a) Liquid-chlorine drum; b) Buffer vessel; c) Flexible coil; d) Chlorine vaporizer; e) Protective membrane; f) Relief safety
valve; g) Rupture disk; h) Barometric leg; i) Water pump; j) Water heater
the risk of concentrating nitrogen trichloride and
other nonvolatile residues in the liquid phase
within the tank. In all other circumstances, the
liquid chlorine should be transferred into a fixed
storage vessel and then vaporized in a special
installation.
Liquid chlorine is discharged by putting the
tank under pressure with dry inert gas or dry
chlorine gas. The inert transfer gas must have a
dew point below − 40 ◦ C at atmospheric pressure and must be clean and free of impurities
such as dust or oil. Before closing the valves,
the tanks must be vented to avoid the risk of high
pressure in the container on account of the additional partial pressure of the inert gas. The use of
an inert gas requires the availability of a chlorine
absorption or neutralization system. Discharge
with pressurized chlorine gas requires a chlorine vaporizer or a special chlorine compressor.
Articulated arms, flexible hoses, and steel coils
are used for the flexible connections. Remotecontrol valves installed close to the ends of the
flexible connections limit leakage in the event of
a failure.
Recommendations on technical equipment,
installation, taking into operation, checks, and
handling are provided by The Chlorine Institute
and Euro Chlor.
12.4. Chlorine Vaporization
When large amounts of chlorine gas are required
or when the chlorination process needs pressurized gas, liquid chlorine must be vaporized and
superheated to avoid liquefaction [242, 243]. It
is advisable to operate the vaporizer at a sufficiently high temperature to accelerate the decomposition of nitrogen trichloride. As a source
of heat, steam with a maximum allowed temperature of 120 ◦ C is used when the vaporizing
system is constructed of mild steel. Water above
60 ◦ C is also suitable, as shown in Figure 71. Direct electrical heating is not appropriate because
there is always a risk of overheating the steel.
Coil-in-bath vaporizers use a coiled tube or
a spiral located in a vessel of hot water (Fig.
71). Generally, they are used for small throughputs; they are simple in design and construction.
Double-envelope vaporizers have compact construction and are easy to operate and to maintain.
Vertical tube vaporizers have a large surface area
and allow a high flow rate. Kettle vaporizers are
also constructed for large unit capacities [244].
Every effort must be taken to avoid the reverse
suction of water or organic materials into the vaporizer. The recommended water and nitrogen
trichloride content of introduced liquid chlorine
must not be exceeded. Vaporizers operating at
low temperature or with a constant liquid level
80
Chlorine
Figure 72. Absorption equipment for the treatment of gases containing chlorine
a) Buffer vessel; b) Vent fan; c) Packed tower; d) Circulating pump; e) Heat exchanger (cooler)
need to be purged to avoid dangerous concentration of nitrogen trichloride [245, 246].
12.5. Treatment of Gaseous Effluents
Gaseous effluents containing chlorine arise from
various sources and must be treated in such a way
as to obtain a tolerable concentration of chlorine
when they are released into the air. The vent gas
may contain other substances, such as hydrogen, organic compounds, CO2 , etc., which must
be considered in design and operation of an effluent treatment installation [247, 248].
Operation of the collection system below atmospheric pressure facilitates the purging of
chlorine vessels, pipes, etc. The risk of corrosion
in dry chlorine installations by moisture from the
treatment system must be excluded. The most
commonly used and recommended reagent is
caustic soda. The effluents are treated in an absorption system, such as packed absorption towers, venturi scrubbers, etc. An example of a flow
sheet for a large plant is shown in Figure 72.
To avoid any formation of solid salts, the
recommended concentration of caustic soda is
< 22 wt %. The operating temperature should
not exceed 55 ◦ C; under normal conditions a
temperature of ca. 45 ◦ C is usual. A cooling
system may be necessary. In large chlorine absorption units, the sodium hypochlorite solution
that is produced can be used in other processes.
Where this is not possible, several methods can
be used to decompose the hypochlorite: controlled thermal decomposition, catalytic decomposition [249], acidification, for example, with
sulfuric acid
NaCl+NaOCl+H2 SO4 → Na2 SO4 +Cl2 +H2 O
12.6. Materials
The choice of material [250] depends on the design and operating conditions and must take into
account all circumstances. A chlorine manufacturer should be consulted to confirm the suitability of a material. Any use of silicone materials
in chlorine equipment should be avoided.
Dry Chlorine Gas (water < 40 ppm by
weight). Carbon steel is the material most used
for dry chlorine gas. It is protected by a thin
layer of ferric chloride. For practical purposes
the recommended temperature of these materials is ≤ 120 ◦ C. High-surface areas, such as
steel wool, or the presence of rust and organic
substances increase the risk of ignition of steel.
The resistance of stainless steels to chlorine
at high temperature increases with the content of
nickel. For stainless steels containing less than
10 wt % nickel, the upper temperature limit is
Chlorine
81
150 ◦ C. High-nickel alloys, such as Monel, Inconel, or Hasteloy C, are suitable up to 350 –
500 ◦ C. Poor mechanical strength limits the use
of nickel. Copper is used for flexible connections
and coils, but it becomes brittle when stressed
frequently.
Because titanium ignites spontaneously in
dry chlorine, it must be avoided. Graphite, glass,
and glazed porcelain are used where there is
a risk of moisture in the dry chlorine gas, and
poly(vinyl chloride) (PVC) or chlorinated PVC
and polyester resins are suitable if the temperature limits of these materials are regarded.
are recommended [253]. In wet chlorine gas,
rubber or synthetic elastomers are acceptable.
Even at temperatures up to 200 ◦ C, PTFE is resistant against wet and dry chlorine gas and liquid chlorine.
Materials resistant because of protection by a
chloride surface layer are not recommended for
protective membranes, rupture disks, and bellows. Suitable materials are tantalum, Hasteloy
C, PTFE, PVDF, Monel, and nickel.
Liquid Chlorine. Unalloyed carbon steel
and cast steel are used with liquid chlorine. Lowtemperature chlorine systems apply fine-grain
steels with a limited tensile strength to guarantee
good conditions for welding. To avoid erosion of
the protective layer, practice is to limit the velocity of the liquid to less than 2 m/s. Organic
materials – rubber lining, ebonites, polyethylene, polypropylene, PVC, chlorinated PVC,
polyester resins, and silicone – are dangerous
[251]. Zinc, tin, aluminum, and titanium are not
acceptable. For certain equipment, copper, silver, lead, and tantalum are appropriate.
In hazard and risk assessment studies, the design
of chlorine installations and equipment and the
operating and maintenance concepts are examined in detail to minimize risks [254]. However,
there remains a certain risk, and all efforts must
be taken to protect people and the environment
in the case of a chlorine emergency. The penetrating odor and the yellow-green color of a
cloud indicate chlorine in the air. If around-theclock surveillance by operators is not possible,
automatic leak detectors are available. Safety in
handling chlorine depends largely on the education and training of employees. An emergency
plan should be brought to the attention of the
personnel involved. Computer-assisted systems
can be used in certain circumstances [255]. Periodic exercises and safety drills should be carried
out.
All people on a chlorine plant are advised to
carry escape-type respirators. The use of filter
masks is prohibited where there is a risk of a
high concentration of chlorine. Anyone who enters an area with high chlorine concentrations
should be equipped with self-contained breathing apparatus and full protective clothing suitable for dealing with liquid chlorine. Protective
equipment, safety showers, eye-wash facilities,
and emergency kits [24] must be quickly accessible.
A means of indicating the actual wind direction should be located near the chlorine installation.
Fixed or mobile water curtains can be used
to divert the dispersion of a chlorine gas cloud
[256]. However, the direct discharge of water
into liquid chlorine and on the area of a chlorine
leak must be avoided.
Wet Chlorine Gas. Wet chlorine gas rapidly
attacks most common metallic materials with
the exception of tantalum and titanium. To assure a protective oxide layer on the surface of
the titanium, sufficient water must be present in
the chlorine gas. If the system does not remain
sufficiently wet, titanium ignites spontaneously
[252].
Most organic materials are slowly attacked
by wet chlorine gas. Rubber-lined iron is successfully used up to 100 ◦ C. At low pressure
and temperature the use of plastic materials like
PVC, chlorinated PVC, and reinforced polyester
resins is advantageous. Polytetrafluoroethylene (PTFE), poly(vinylidene fluoride) (PVDF),
and fluorinated copolymers like tetrafluoroethylene – hexafluoropropylene (FEP) are resistant
even at higher temperature. Ceramics have been
progressively replaced by plastics. Impregnated
graphite is suitable up to 80 ◦ C; the impregnation should be resistant to wet chlorine.
Materials for Special Parts. After the ban
of asbestos as material for gaskets, substitutes
12.7. Safety
82
Chlorine
In most countries, chlorine manufacturers
have organized groups of experts who are well
versed and drilled in handling chlorine and can
be called at any time in case of chlorine emergency.
The Chlorine Institute has released pamphlets
and recommendations covering all aspects of
safety, e.g., first aid [257], emergency response
plans [258], protective equipment [259], prevention of injuries to personnel [260], prevention of
chlorine releases [261], and estimating the area
affected by a chlorine release [262]. Emergency
kits have been developed for sealing leaks in
chlorine containers, drums, and tank cars [263].
Euro Chlor offers a Chlorine Safety Manual
[264], recommendations for emergency intervention [265], and for safe design, construction,
operation of equipment [266].
Emergency plans have been established for
accidents during transportation and use, e.g.,
CHLOREP [267] in North America and TUIS
(Transport Unfall Informations System) in Germany.
13. Quality Specifications and
Analytical Methods
13.1. Quality Specifications
Liquid chlorine of commercial quality must have
a purity of at least 99.5 wt % [268]. The water content is < 0.005 wt %, and solid residues
are < 0.02 wt %. The impurities are mainly CO2
(≤ 0.5 wt %), N2 , and O2 (each 0.1 – 0.2 wt %).
There are traces of chlorinated hydrocarbons
(originating from rubberized or plastic piping)
and inorganic salts such as ferric chloride. The
chlorine may also contain small amounts of
bromine or iodine, depending on the purity of
the salt used in the electrolytic process.
13.2. Analytical Methods
Industrial liquid chlorine is mainly analyzed by
the methods in ISO regulations. The liquid chlorine is evaporated at 20 ◦ C, and this gas is then
analyzed.
Sampling
Moisture
Chlorine content
Gaseous components
NCl3
Mercury
ISO 1552 [269, 270]
ISO 2121 [271], ASTM E410
[272]
ISO 2202 [273]
ISO 2120 [274], ASTM E412
[275]
DIN 38 408, part 4 [276]
[277]
[278], ASTM E506 [279]
The residue is weighed, and the organic constituents are taken up in acetone, hexane, or diethyl ether and determined by gas chromatography. The inorganic residue is analyzed. For
quick analysis, liquid chlorine can be introduced
directly onto a silica gel column of a gas chromatograph.
Chlorine Gas. The chlorine gas can be analyzed for chlorine content, gaseous impurities,
hydrogen, organics, and moisture:
1) Chlorine content. One method for process
monitoring and control of chlorine concentration is measurement of thermal conductivity.
2) Gaseous impurities. A known amount of chlorine gas is passed through a solution of potassium iodide or phenol to absorb the chlorine.
The residual gases (O2 , N2 , H2 , CO, CO2 )
are collected in a gas burette, measured, and
analyzed by gas chromatography or with an
Orsat apparatus.
3) Hydrogen. A known amount of air is added to
the residual gas after removal of the chlorine
to ensure excess oxygen, and the volume reduction is measured after the hydrogen is consumed on a heated platinum coil. The hydrogen content can be continuously monitored by
thermal conductivity measurement.
4) Organics. Organic components can be determined most conveniently by gas chromatography.
5) Moisture. A known amount of chlorine gas
is passed through a drying tube filled with
a weighed amount of phosphorus pentoxide.
The moisture content is determined from the
weight gain of the drying tube (ISO 2121).
Continuous determination can be carried out,
e.g., by absorption with phosphorus pentoxide and measuring the current required to electrolyze the absorbed water or by the electrical
conductivity after absorption in sulfuric acid.
Chlorine
Detection of Chlorine. Chlorine can be recognized by smell or color. Small amounts can
be detected by the blue coloration of starch –
iodide paper, although other oxidizing agents
can produce the same effect. Another method for
chlorine detection depends on its ability to combine with mercury. If the unknown gas mixture is
shaken with water and mercury, all the chlorine
disappears and the remaining water has a neutral
reaction. However, if the chlorine contains some
hydrogen chloride, the water becomes acidic and
reacts with silver nitrate solution to give a white
precipitate (AgCl) that is soluble in aqueous ammonia. Leaks in pipes or equipment are detected
by testing with the vapor from aqueous ammonia: a thick white cloud of chloride forms.
Quantitative Determination of Free Chlorine. The gas mixture can be shaken with a
potassium iodide solution, and the liberated iodine can be then determined by titration. Chlorine in alkaline solution can be reduced to
chloride by potassium or sodium arsenite, and
the arsenite can be then oxidized to arsenate.
The end point is detected by spot tests with
starch – iodide paper. Excess arsenite is backtitrated with acidified potassium bromate solution. Small amounts of chlorine, e.g., in drinking
water, can be determined by photometric measurement of the yellow color produced by the
reaction with o-tolidine in hydrochloric acid solution [280].
To determine both chlorine and carbon dioxide, the chlorine is absorbed by a solution that
contains known amounts of acid and potassium
arsenite, and the chlorine is determined by backtitration of the arsenite. The carbon dioxide,
which is not absorbed by this solution, is then
absorbed by potassium hydroxide solution.
Detection Tubes. Commercial detection
tubes (Drägerwerk, Lübeck; Auergesellschaft,
Berlin) are available for measuring chlorine
in air. They have various ranges: 0.2 – 3, 2 –
30, and 50 – 500 ppm. The Chlorometer (ZeissIkon, Berlin) can determine the free chlorine
content of water in a few minutes.
For the protection of the environment and
control of working conditions, traces of chlorine as small as 0.01 – 10 ppm must be determined. There are many types of apparatus on
the market for measuring workplace concentrations or emissions. They depend on physicochemical methods, such as conductometry, gal-
83
vanometry, potentiometry, colorimetry, and UV
spectroscopy [281].
14. Uses
The first industrial use of chlorine was to produce bleaching agents for textiles and paper and
for cleaning and disinfecting. These were liquid bleaches (solutions of sodium, potassium,
or calcium hypochlorite) or bleaching powder
(chlorinated lime). Chlorine was then regarded
merely as a useful chemical agent.
Since 1900s, chlorine has achieved constantly increasing importance as a raw material
for synthetic organic chemistry. Chlorine is an
essential component of a multitude of end products, which are used as materials of construction, solvents, pesticides, etc. In addition, it is
contained in intermediates that are used to make
chlorine-free end products. It is these areas of
use that allow chlorine production to increase.
The direct use of elemental chlorine, e.g., for
sterilizing water, has declined in some areas,
but not in others. For example, in Germany, it
is < 0.1 %, but in the United States it is 6 %
(1982). The WHO recommends chlorine as the
most useful agent for water disinfection [282].
The percentage of world chlorine consumption of various product groups in 1997 was as
follows [283]:
Vinyl chloride
Misc. organic products
Solvents
Pulp and paper
Water treatment
Others (inorganic products, etc)
33
6
19
5
6
31
The number of possible reactions of chlorine, and therefore the number of intermediates
and end products, is remarkably large. Some important reactions are shown in Figure 73 along
with the areas of application of the end products
[284]. Figure 74 also shows some of the applications of chlorine. About one-third of the finished
products are chlorine-free, although made with
the help of chlorine [285].
Inorganic Products. The use of chlorine
to produce pure hydrochloric acid, bleaching
agents, and other inorganic products is increasing, although slowly enough that this forms a
84
Chlorine
Figure 73. Important reactions of chlorine and the uses of the end products
steadily decreasing proportion of the total consumption. The production of some metal chlorides is increasing, e.g., titanium tetrachloride,
aluminum chloride, and silicon tetrachloride.
States: 13 % in 1983, 9 % in 1994, 6 % in 2000,
0 % in 2010 (forecast); the decline in Canada
and in Scandinavia is even sharper (1983: 60 –
70 %; 1998: near 0 %) [286].
Bleaching. Countries with large paper and
cellulose industries had a large chlorine consumption for bleaching purposes. Ecological
concerns about AOX (absorbable organic halogens) in the effluents of papermills have led to replacement of elemental chlorine by other bleaching agents, including ozone, hydrogen peroxide, and chlorine dioxide. As a consequence, the
use of chlorine is dropping sharply; e.g., United
Organic Products. Organic intermediates
and end products accounted for up to 85 % of the
chlorine use in developed countries. The product with the largest chlorine requirement is vinyl
chloride, which accounted for 33 % of world
chlorine consumption in 1997. This proportion
will increase further and outweigh consumption
losses in other areas and is therefore the reason
for increasing chlorine capacities worldwide.
Chlorine
85
Figure 74. The chlorine tree (courtesy of Euro Chlor)
Chlorine reacts with methane to give the
solvents methyl chloride, methylene chloride,
chloroform, and carbon tetrachloride. Butane
and pentane are the starting materials for
hexachlorobutadiene and hexachloropentadiene, the basis of many pesticides. Chlorinated
paraffins are starting materials for detergent raw
materials, plasticizers, and lubrication additives.
The chlorination of ethylene first gives 1,2dichloroethane, which is an important starting
material for a variety of further products, including vinyl chloride, ethylene oxide, ethylene glycol, ethylenediamine, and chloral. Other
products of the chlorination of ethylene are
trichloroethylene, perchloroethylene, etc.
The following products are obtained from
propylene: propylene oxide, propylene glycol,
carbon tetrachloride (via propylene dichloride),
glycerine (via allyl chloride), and epoxy resins
and synthetic rubber (via epichlorhydrin).
Aromatic derivatives are also important.
The reaction of chlorine with benzene produces the following compounds, among others: monochlorobenzene, dichlorobenzene, and
hexachlorobenzene, which are used as solvents or pesticides. A more important use of
monochlorobenzene, however, is as an intermediate in the manufacture of aniline, and dyes. A
further important synthesis route is to start from
toluene to produce toluenediamine and with the
86
Chlorine
help of phosgene to convert this to toluene diisocyanate, which is used in plastics.
Environmental Aspects. The production
and use of chlorine, as well as certain chlorinated substances bear risks for human health and
the environment. International discussions are
underway on legally-binding conventions and
protocols thereof. Matters of concern include
the use of mercury, asbestos, chlorine containing materials which are persistant, toxic, and
liable to bioaccumulate (e.g., polychlorinated
biphenyls, dioxins and furans, DDT, endrine,
chlordane), and chemicals having an ozonedepletion potential (e.g., chlorofluorocarbons).
Regarding the principles of “Responsible Care”
and of “Sustainable Development” [287] the
chemical industry is managing this challenge:
Once a hazard is realised, the substance in question is handled individually by investigating its
properties and its impact on the environment on
a scientific basis. Risk assessments and risk –
benefit studies are prepared, the results are published as fact sheets, position papers, scientific
publications and contributions to the discussion
with governments, authorities, and the public.
Practical measures for reducing the risks are
continuous efforts to reduce or eliminate emissions (mercury, asbestos), offer recycling concepts for products (PVC, chlorinated solvents),
improve production processes and transportation safety, cease the production (CFCs, pentachlorophenol in most countries) or marketing
of problematic compoundes, and developing
for alternative processes (bleaching in the pulp
and paper industry). All these measures are influencing the use pattern of chlorine: the use
for bleaching, for C1 compounds, for solvents,
and for tetraethyllead will further decrease; the
importance of organic chemicals will rise and
PVC will expand especially in the construction
sector. A study conducted by ECOTEC [288]
illustrates the chlorine flow within the industrial
production system and to external consumers.
15. Economic Aspects
In Western Europe in 1995 [289]:
– Almost 2 × 106 jobs were related to chlorine
– 55 % of European chemical turnover depended on chlorine
– 85 % of pharmaceuticals are made using chlorine
– 98 % of Western Europe’s drinking water is
purified by chlorination
These figures demonstrate the economic importance of the chlorine industry.
Figure 75 shows how world chlorine production has developed from 1965 to 1995. The figures are only estimates, as many important nations do not publish this information.
Figure 75. World chlorine production (1965 – 1995)
Chlorine capacities 1995 (in 103 t/a) in the
main producer countries was as follows [290]:
United States
Japan
Former Soviet Union
China
Germany
Brazil
Canada
France
United Kingdom
India
Italy
11 860
4250
3800
3750
3690
1660
1630
1580
1270
1230
1130
World chlorine capacity was about
40 × 106 t/a in 1983, it rose to 48 × 106 t/a in
1998, and is forecast to reach 55 × 106 t/a by
2005 [283].
Figure 76 demonstrates the growth of chlorine capacity by economic regions [291]. Compared to 1983, the capacities in 1997 remained
nearly constant in Western Europe, Eastern Europe, and Africa. In North America, reductions
Chlorine
87
Figure 76. Chlorine production capacities by economic regions (1983 and 1997)
in Canada were outweighed by growth in the
United States. Huge capacities were erected in
the Middle East and above all in Asia and Oceania (China, India), where capacities were tripled.
From the growth between 1997 and 2005, North
America’s share will be 35 %, Asia’s 40 %, and
the Middle East’s 15 % [283]. After 2000 China
will become the second biggest producer after
the United States, with an annual chlorine capacity of about 5.5 × 106 t.
All chlorine producers are listed in [292]. A
detailed review of the chlor-alkali market for
1994 is given in [293].
The production costs of 1 electrochemical
∧
unit (= 1 t chlorine + 1.13 t sodium hydroxide)
depends up to 60 % on the price of electricity. At
3.5 c/ /kWh, they are about 250 $/ecu (Fig. 77), so
the total value of production is ca 24 × 109 $/a.
The chlor-alkali industry is one of the biggest
consumers of electrical energy, consuming ca.
0.15 × 1012 kWh/a.
A prediction of the market prices for chlorine
and for caustic soda is very difficult. Firstly, most
chlorine is used captively to avoid transportation, or it is sold on the basis of long-term contracts. In relation to production, the amounts on
the spot markets are small. Secondly, the production of chlorine is strictly coupled to that of the
caustic soda. A strong demand for chlorine creates a surplus of caustic, leading to high prices
for the chlorine and a drop in prices for caustic;
weak demand for chlorine reduces the available
amount of caustic, leading to high caustic prices
and low chlorine prices. This situation seems to
change in cycles every 6 to 8 years. For example,
from 1988 to 1991 the market value for chlorine fell from 150 $/t down to 0, but from March
1992 to August 1993, the price rose to 230 $/t.
The price for caustic was 320 $/t in 1991; it fell
to 50 $/t in March 1994 and climbed to 270 $/t
in December 1994 (Fig. 77) [294].
Because of the fixed ratio of chlorine to caustic in production and the different location of
uses, worldwide trade flows have been created
for both products. Chlorine is traded as chlorinated derivatives, particularly as EDC, VCM,
and PVC, accounting for 85 %, and chlorinated
solvents. The equivalent annual amount of chlorine is approximately 4.5 × 106 t (1995 to 2000).
The main exporting countries are the United
States (to South East Asia, South America), Russia (to West Europe), West Europe (to South East
Asia, China), Japan (to China), and Middle East
Asia (to Japan , China). [295].
16. Toxicology
Chlorine gas is dangerous to health because it
is a powerful oxidant. In the physiological pH
range, it is converted to hypochlorous acid, a cytotoxic substance. The extent to which the cells
are damaged depends on the gas concentration,
the exposure time, the water content of the tis-
88
Chlorine
Figure 77. Production costs and product prices in the United states from 1987 to 1997 [294]
a) Product value for 1 ecu (1 t Cl2 + 1.13 t NaOH; b) Price for caustic soda; c) Production costs for 1 ecu at 3.5 c//kWh electricity
price; d) Price for chlorine into PVC
sue, and the health of the person exposed to the
gas. Besides getting into the eyes, larynx, and
trachea, chlorine also reaches the bronchi and
the bronchioles. Because of its moderate solubility in water, chlorine affects the alveoli only
at high concentrations. In this respect, chlorine
differs from other gases with low water solubility and high lipid solubility, such as phosgene,
nitrogen monoxide, and nitrogen dioxide. Initial
moderate bronchial irritation is followed by the
development of a toxic pulmonary edema because of increased alveolar injury.
The olfactory threshold of chlorine gas is
0.2 – 3.5 mL/m3 . Prolonged exposure seems to
raise the olfactory threshold. Concentrations of
3 – 5 mL/m3 are tolerated for up to 30 min without any subjective feeling of malaise. At concentrations between 5 and 8 mL/m3 , mild irritation
of the upper respiratory tract and the conjunctiva is observed. In addition, running of the eyes
and coughing are observed at concentrations of
15 mL/m3 and higher. Above 30 mL/m3 , the following symptoms are observed: nausea, vomiting, oppressive feeling, shortness of breath, and
fits of coughing, sometimes leading to bronchial
spasms. Exposure to 40 – 60 mL/m3 leads to the
development of toxic tracheobronchitis. After a
latent period of several hours with fewer symptoms, pulmonary edema may occur because of
alveolar membrane destruction. This is indicated
by increased shortness of breath, restlessness,
and cyanosis. Subsequently, a further complication can occur after several days in the form of
pneumonia caused by superinfection of the injured pulmonary tissue.
A clear dose – effect relationship of chlorine gas at different concentrations in humans
has not yet been published. On the basis of
results obtained from animal experiments, the
LC50 for healthy humans is assumed to be
300 – 400 mL/m3 at a 30-min exposure [296].
No deaths have occurred in animal experiments
at 30-min exposures for concentrations below
50 mL/m3 . Death following acute intoxication
is caused by a fulminant pulmonary edema.
Many investigations deal with the toxicity of low chlorine concentrations. Investigations in humans indicate the possibility of reversible damage to the lung function parameters
at concentrations starting at 1.0 mL/m3 , whereas
0.5 mL/m3 is a no-observable-effect-level [297,
298]. In monkeys no effects have been observed
Chlorine
at 0.5 mL/m3 in a one-year study [299]. Longterm investigations of workers exposed to chlorine, e.g., in chlor-alkali electrolysis plants or in
pulp manufacture, however, do not indicate increased rates of mortality or morbidity caused
by pulmonary diseases [300 – 303].
No indications of carcinogenicity or mutagenicity of chlorine have been detected in animal experiments or encountered in industrial
medicine [304]. In Germany, the MAK value
is 0.5 mL/m3 (1.5 mg/m3 ) with a 15-min excursion factor of 1 [304]; in the USA, the
TLV is 0.5 mL/m3 (1.5 mg/m3 ) with a STEL of
1.0 mL/m3 (3 mg/m3 ).
17. Acknowledgement
The entire topic was coordinated by Peter
Schmittinger
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